Synthesis gas conversion process

ABSTRACT

The disclosed invention relates to a method for restarting a synthesis gas conversion process which has stopped. The synthesis gas conversion process may be conducted in a conventional reactor or a microchannel reactor. The synthesis gas conversion process may comprise a process for converting synthesis gas to methane, methanol or dimethyl ether. The synthesis gas conversion process may be a Fischer-Tropsch process.

This application is a continuation of U.S. patent application Ser. No.16/460,038, filed Jul. 2, 2019, which is a continuation of applicationSer. No. 15/208,902, filed on Jul. 13, 2016, now U.S. Pat. No.10,358,604, which is a continuation of application Ser. No. 15/178,902,filed on Jun. 10, 2016, now abandoned. This application claims priorityto U.S. Provisional Application No. 62/174,772, filed on Jun. 12, 2015.These prior applications are incorporated herein by reference.

TECHNICAL FIELD

This invention relates to a synthesis gas conversion process, and moreparticularly to a method for restarting a synthesis gas conversionprocess that has stopped.

BACKGROUND

Synthesis gas comprises H₂ and CO. Synthesis gas conversion processesinclude processes for converting synthesis gas in the presence of asynthesis gas conversion catalyst to a synthesis gas conversion product.The synthesis gas conversion product may comprise a methane, methanol,or dimethyl ether. The methane that is formed may be referred to assynthetic natural gas. The synthesis gas conversion process may comprisea Fischer-Tropsch process, and the product may comprise aFischer-Tropsch product.

SUMMARY

Synthesis gas conversion production runs are typically conducted overextended periods of time. For example, a Fischer-Tropsch production runmay be conducted over an extended period of time of, for example, atleast about 300 hours up to about 8000 hours or more. In a typicalproduction run the catalyst activity declines over time and iscompensated for by increasing the reaction temperature. However, at somepoint in time the production run will be stopped, either purposely oraccidentally, and the problem addressed with this invention relates toproviding a rapid restart of the production run with no or minimal lossof catalytic activity. A problem to be solved when stopping andrestarting the process relates to avoiding conditions which causecatalyst deactivation when the feed is interrupted and restarted. Thiscan be caused by continued high rates of reaction consuming remainingsynthesis gas or hydrogen in the system and exposing the catalyst toundesirable amounts of carbon monoxide or water. Another problem relatesto achieving a fast restart to maximize production/operating time onstream. Another problem to be solved relates to the fact that typicalsolutions to the above-indicated problems tend to be only effective fora limited time duration after the reaction feed is interrupted. As aresult, it is often necessary to take action to avoid significantcatalyst deactivation. Another problem relates to recovering lostcatalyst activity that typically occurs after limited time duration withthe short catalyst rejuvenation or regeneration procedures.

This invention provides solutions to these problems. With thisinvention, a synthesis gas conversion production run (e.g., aFischer-Tropsch production run) may be restarted and achieve fullcapacity (i.e., the same or substantially the same catalytic activity orCO conversion rate as prior to the stoppage) within, for example, up toabout 48 hours, or up to about 36 hours, or up to about 24 hours, or upto about 12 hours, or up to about 6 hours, or up to about 3 hours, or upto about 2 hours, or up to about 1 hour, or up to about 0.5 hour, or upto about 0.1 hour, or from about 0.01 to about 48 hours, or from about0.01 to about 36 hours, or from about 0.01 to about 24 hours, or fromabout 0.01 to about 12 hours, or from about 0.01 to about 6 hours, orfrom about 0.01 to about 3 hours, or from about 0.01 to about 1 hour, orfrom about 0.01 to about 0.5 hour, or from about 0.1 to about 48 hours,or from about 0.1 to about 36 hours, or from about 0.1 to about 24hours, or from about 0.1 to about 12 hours, or from about 0.1 to about 6hours, or from about 0.1 to about 3 hours, or from about 0.1 to about 2hours, or from about 0.1 to about 1 hour, or from about 0.1 to about 0.5hour, from the time the flow of synthesis gas into the reactor isrestarted. The activity temperature delta for the process may be up toabout 5° C., or up to about 3° C., or up to about 1° C., after restart.The relative activity ratio for a process utilizing this invention maybe at least about 0.85, or at least about 0.92, or at least about 0.97,or at least about 1.00.

This invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas conversion product and flowingeffluent comprising the synthesis gas conversion product out of thereactor, the method comprising: (A) stopping the flow of the synthesisgas into the reactor and the flow of the effluent out of the reactor fora period of time of up to about 48 hours, or up to about 36 hours, or upto about 24 hours, or up to about 12 hours, or up to about 6 hours, orup to about 3 hours, or up to about 1 hour, or up to about 0.5 hour, orup to about 0.1 hour, or from about 0.01 to about 48 hours, or fromabout 0.01 to about 36 hours, or from about 0.01 to about 24 hours, orfrom about 0.01 to about 12 hours, or from about 0.01 to about 6 hours,or from about 0.01 to about 3 hours, or from about 0.01 to about 1 hour,or from about 0.01 to about 0.5 hour, or from about 0.1 to about 48hours, or from about 0.1 to about 36 hours, or from about 0.1 to about24 hours, or from about 0.1 to about 12 hours, or from about 0.1 toabout 6 hours, or from about 0.1 to about 2 hours, or from about 0.1 toabout 3 hours, or from about 0.1 to about 1 hour, or from about 0.1 toabout 0.5 hour; and (B) restarting the flow of synthesis gas into thereactor and the flow of effluent out of the reactor.

This invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas conversion product and flowingeffluent comprising the synthesis gas conversion product out of thereactor, the method comprising: (A) stopping the flow of the synthesisgas into the reactor and the flow of the effluent out of the reactor fora period of time of up to about 48 hours, or up to about 36 hours, or upto about 24 hours, or up to about 12 hours, or up to about 6 hours, orup to about 3 hours, or up to about 1 hour, or from about 0.01 to about48 hours, or from about 0.01 to about 36 hours, or from about 0.01 toabout 24 hours, or from about 0.01 to about 12 hours, or from about 0.01to about 6 hours, or from about 0.01 to about 3 hours, or from about0.01 to about 1 hour, or from about 0.01 to about 0.5 hour, or fromabout 0.1 to about 48 hours, or from about 0.1 to about 36 hours, orfrom about 0.1 to about 24 hours, or from about 0.1 to about 12 hours,or from about 0.1 to about 6 hours, or from about 0.1 to about 3 hours,or from about 0.1 to about 2 hours, or from about 0.1 to about 1 hour,or from about 0.1 to about 0.5 hour; wherein prior to step (A) thepressure within the reactor is at a pre-stoppage pressure and duringstep (A) the pressure within the reactor is reduced to a level lowerthan the pre-stoppage pressure; (B) restoring the pressure within thereactor to the pre-stoppage pressure; and (C) restarting the flow of thesynthesis gas into the reactor and the flow of effluent out of thereactor.

This invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas product and flowing effluentcomprising the synthesis gas conversion product out of the reactor, themethod comprising: (A) stopping the flow of the synthesis gas into thereactor and the flow of the effluent out of the reactor for a period oftime of up to about 48 hours, or up to about 36 hours, or up to about 24hours, or up to about 12 hours, or up to about 6 hours, or up to about 3hours, or up to about 1 hour, or from about 0.01 to about 48 hours, orfrom about 0.01 to about 36 hours, or from about 0.01 to about 24 hours,or from about 0.01 to about 12 hours, or from about 0.01 to about 6hours, or from about 0.01 to about 2 hours, or from about 0.01 to about1 hour, or from about 0.01 to about 0.5 hour, or from about 0.1 to about48 hours, or from about 0.1 to about 36 hours, or from about 0.1 toabout 24 hours, or from about 0.1 to about 12 hours, or from about 0.1to about 6 hours, or from about 0.1 to about 3 hours, or from about 0.1to about 2 hours, or from about 0.1 to about 1 hour, or from about 0.1to about 0.5 hour; (B) flowing hydrogen into the reactor and restartingthe flow of effluent out of the reactor to purge the reactor and torejuvenate the catalyst; and (C) restarting the flow of synthesis gasinto the reactor and the flow of effluent out of the reactor. In anembodiment, the desired reaction temperature is in the range from about150° C. to about 300° C. and during step (B) the temperature within thereactor is increased to a temperature above the desired reactiontemperature. In an embodiment, during step (B) the temperature withinthe reactor is increased to a temperature of up to about 350° C., or upto about 400° C. In an embodiment, the reactor is cooled during step (B)to a temperature in the range from about 150° C. to about 200° C., orabout 170° C. In an embodiment, prior to step (A) the temperature withinthe reactor is at a pre-stoppage temperature, and subsequent to step (B)the temperature within the reactor is reduced to a level below thepre-stoppage temperature. In an embodiment, prior to step (A) thepressure within the reactor is at a pre-stoppage pressure, andsubsequent to step (B) the pressure within the reactor is reduced to alevel below the pre-stoppage pressure.

In an embodiment, the temperature of the reactor is controlled with aheat exchange fluid flowing in a heat exchanger in thermal contact withthe reactor.

In an embodiment, the purge may be at a flow rate lower than or equal tothe synthesis gas feed rate to the reactor prior to stoppage. In anembodiment, the purge may be at a flow rate equal to or higher than thesynthesis gas feed rate to the reactor prior to stoppage.

This invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas conversion product and flowingeffluent comprising the synthesis gas conversion product out of thereactor, the method comprising: (A) stopping the flow of synthesis gasinto the reactor and the flow of effluent out of the reactor for aperiod of time of up to about 48 hours, or up to about 36 hours, or upto about 24 hours, or up to about 12 hours, or up to about 6 hours, orup to about 3 hours, or up to about 1 hour, or from about 0.01 to about48 hours, or from about 0.01 to about 36 hours, or from about 0.01 toabout 24 hours, or from about 0.01 to about 12 hours, or from about 0.01to about 6 hours, or from about 0.01 to about 3 hours, or from about0.01 to about 1 hour, or from about 0.01 to about 0.5 hour, or fromabout 0.1 to about 48 hours, or from about 0.1 to about 36 hours, orfrom about 0.1 to about 24 hours, or from about 0.1 to about 12 hours,or from about 0.1 to about 6 hours, or from about 0.1 to about 3 hours,or from about 0.1 to about 2 hours, or from about 0.1 to about 1 hour,or from about 0.1 to about 0.5 hour; (B) flowing natural gas into thereactor and restarting the flow of effluent out of the reactor to purgethe reactor; and (C) restarting the flow of synthesis gas into thereactor. In an embodiment, prior to step (A) the temperature within thereactor is at a pre-stoppage temperature, and subsequent to step (B) thetemperature within the reactor is reduced to a level below thepre-stoppage temperature. In an embodiment, prior to step (A) thepressure within the reactor is at a pre-stoppage pressure, andsubsequent to step (B) the pressure within the reactor is reduced to alevel below the pre-stoppage pressure. The natural gas may be adesulfurized natural gas. In an embodiment, the purge may be at a flowrate lower than or equal to the synthesis gas feed rate to the reactorprior to stoppage. In an embodiment, the purge may be at a flow rateequal to or higher than the synthesis gas feed rate to the reactor priorto stoppage.

This invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas product and flowing effluentcomprising the synthesis gas conversion product out of the reactor, themethod comprising: (A) stopping the flow of synthesis gas into thereactor; (B) flowing hydrogen into the reactor to purge the reactor; and(C) restarting the flow of synthesis gas into the reactor.

This invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas conversion product and flowingeffluent comprising the synthesis gas conversion product out of thereactor, the method comprising: (A) stopping the flow of synthesis gasinto the reactor; (B) flowing natural gas into the reactor to purge thereactor; and (C) restarting the flow of synthesis gas into the reactor.

This invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas conversion product and flowingeffluent comprising the synthesis gas conversion product out of thereactor, the method comprising: (A) stopping the flow of synthesis gasinto the reactor and the flow of effluent out of the reactor for aperiod of time of up to about 48 hours, or up to about 36 hours, or upto about 24 hours, or up to about 12 hours, or up to about 6 hours, orup to about 3 hours, or up to about 1 hour, or from about 0.01 to about48 hours, or from about 0.01 to about 36 hours, or from about 0.01 toabout 24 hours, or from about 0.01 to about 12 hours, or from about 0.01to about 6 hours, or from about 0.01 to about 3 hours, or from about0.01 to about 1 hour, or from about 0.01 to about 0.5 hour, or fromabout 0.1 to about 48 hours, or from about 0.1 to about 36 hours, orfrom about 0.1 to about 24 hours, or from about 0.1 to about 12 hours,or from about 0.1 to about 6 hours, or from about 0.1 to about 3 hours,or from about 0.1 to about 2 hours, or from about 0.1 to about 1 hour,or from about 0.1 to about 0.5 hour; and (B) restarting the flow ofsynthesis gas into the reactor and the flow of the effluent out of thereactor, the temperature of the synthesis gas flowing into the reactorbeing within about 10° C., or about 5° C., or about 2° C., or about 1°C. of the desired reaction temperature. In an embodiment, thetemperature of the synthesis gas flowing into the reactor during step(B) is at about the desired reaction temperature. In an embodiment, thedesired reaction temperature is controlled with a heat exchange fluidflowing in a heat exchanger in thermal contact with the reactor, thetemperature of the heat exchange fluid in the heat exchanger being up toabout 10° C., or up to about 5° C., or up to about 2° C., or up to about1° C. lower than the desired reaction temperature. In an embodiment, thesynthesis gas conversion catalyst comprises a wet catalyst.

This invention relates to a method for restarting a Fischer-Tropschprocess wherein the Fischer-Tropsch process comprises flowing synthesisgas into a reactor in contact with a Fischer-Tropsch catalyst at adesired reaction temperature and pressure to produce a Fischer-Tropschproduct, the method comprising: (A) stopping the flow of synthesis gasinto the reactor and the flow of effluent out of the reactor for aperiod of time of up to about 48 hours, or up to about 36 hours, or upto about 24 hours, or up to about 12 hours, or up to about 6 hours, orup to about 3 hours, or up to about 1 hour, or from about 0.01 to about48 hours, or from about 0.01 to about 36 hours, or from about 0.01 toabout 24 hours, or from about 0.01 to about 12 hours, or from about 0.01to about 6 hours, or from about 0.01 to about 3 hours, or from about0.01 to about 1 hour, or from about 0.01 to about 0.5 hour, or fromabout 0.1 to about 48 hours, or from about 0.1 to about 36 hours, orfrom about 0.1 to about 24 hours, or from about 0.1 to about 12 hours,or from about 0.1 to about 6 hours, or from about 0.1 to about 3 hours,or from about 0.1 to about 2 hours, or from about 0.1 to about 1 hour,or from about 0.1 to about 0.5 hour; and (B) flowing hydrogen into thereactor in contact with the catalyst at a temperature of up to about400° C. and restarting the flow of effluent out of the reactor; (C)flowing air into the reactor in contact with the Fischer-Tropschcatalyst at a temperature in the range from about 70° C. to about 350°C., or about 100° C. to about 350° C., or about 150° C. to about 350°C., or about 200° C. to about 350° C., or about 250° C. to about 300°C., for a period of time in the range from about 1 to about 24 hours, orabout 1 to about 12 hours; (D) flowing hydrogen into the reactor incontact with the catalyst at a temperature of up to about 400° C. toregenerate the catalyst; and (E) restarting the flow of synthesis gasinto the reactor.

This invention relates to a method of restarting a synthesis gasconversion process wherein the process is conducted in a plantcomprising a plurality of reaction trains, each reaction traincomprising a synthesis gas conversion reactor containing a synthesis gasconversion catalyst, the reaction trains being connected to a reactantfeed stream comprising fresh synthesis gas, the method comprising: (A)flowing the reactant feed stream at an overall process flow rate to theplurality of reaction trains in the plant; (B) dividing the reactantfeed stream into a plurality of reactant substreams; (C) flowing eachreactant substream through a separate reaction train to convert thereactants in the reactant substream to a synthesis gas conversionproduct; (D) stopping the flow of a reactant substream to one of thereaction trains; and (E) continuing to flow the reactant feed stream tothe remainder of reaction trains in the plant to provide a flow rate offresh synthesis gas to the plant that is the same or substantially thesame (i.e., within up to about 15%, or up to about 10%, or up to about5%, or up to about 2%, or up to about 1%) as the flow rate of freshsynthesis gas in step (A). That is, even though one or more reactiontrains may be taken off line, the overall flow rate of fresh synthesisgas to the plant may remain the same or substantially the same. Anydecrease in overall capacity may be less than the capacity that has beentaken off line during step (D). Thus, in an embodiment, during step (E)the overall process flow rate of fresh synthesis gas to the plant may bethe same as the overall process flow rate of fresh synthesis gas used instep (A). In an embodiment, the flow of effluent from the one of thereaction trains is stopped during step (D). In an embodiment, duringstep (C) a mixture of fresh synthesis gas and a recycled tail gas mayflow into the reactor of one or more of the reaction trains. In anembodiment, during step (E) the flow of recycled tail gas into thereactor of the one or more of the remainder of the reaction trains inthe plant may be stopped.

This invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas conversion product and flowingeffluent comprising the gas conversion product out of the reactor,wherein the reactor contains water vapor, the method comprising: (A)stopping the flow of synthesis gas into the reactor and the flow ofeffluent out of the reactor; (B) flowing hydrogen into the reactor incontact with the catalyst and/or removing water vapor from the reactor;and (C) restarting the flow of the synthesis gas into the reactor andthe flow of the effluent out of the reactor. During step (B) thehydrogen may be selectively introduced into the reactor without purgingof the synthesis gas in contact with the catalyst. This may avoidextinction of hydrogen leading to carbon monoxide being present in theabsence of hydrogen, leading to carbonaceous deposits on the catalystand therefore activity loss. During step (C) the water vapor may beselectively removed from the process gas mixture in the reactor. Thismay avoid the risk of catalyst deactivation from water condensation orsintering in presence of steam. One way of accomplishing removal ofwater vapor would be to provide a dessicant in a vessel or piping influid communication with the reactor at a similar pressure.

This invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas conversion product and flowingeffluent comprising the gas conversion product out of the reactor, themethod comprising: (A) stopping the flow of synthesis gas into thereactor and the flow of effluent out of the reactor for a period of timeof up to about 48 hours, or up to about 36 hours, or up to about 24hours, or up to about 12 hours, or up to about 6 hours, or up to about 3hours, or up to about 1 hour, or from about 0.01 to about 48 hours, orfrom about 0.01 to about 36 hours, or from about 0.01 to about 24 hours,or from about 0.01 to about 12 hours, or from about 0.01 to about 6hours, or from about 0.01 to about 3 hours, or from about 0.01 to about1 hour, or from about 0.01 to about 0.5 hour, or from about 0.1 to about48 hours, or from about 0.1 to about 36 hours, or from about 0.1 toabout 24 hours, or from about 0.1 to about 12 hours, or from about 0.1to about 6 hours, or from about 0.1 to about 3 hours, or from about 0.1to about 2 hours, or from about 0.1 to about 1 hour, or from about 0.1to about 0.5 hour; (B) flowing nitrogen gas into the reactor to purgethe reactor; and (C) restarting the flow of the synthesis gas into thereactor and the flow of the effluent out of the reactor. In anembodiment, prior to step (A) the temperature within the reactor is at apre-stoppage temperature, and subsequent to step (B) the temperature isreduced to a level below the pre-stoppage temperature. In an embodiment,prior to step (A) the pressure within the reactor is at a pre-stoppagepressure, and subsequent to step (B) the pressure is reduced to a levelbelow the pre-stoppage pressure.

This invention relates to a method for restarting a Fischer-Tropschprocess, wherein the Fischer-Tropsch process comprises flowing synthesisgas into a reactor in contact with a Fischer-Tropsch catalyst at adesired reaction temperature and pressure to produce a Fischer-Tropschproduct and flowing effluent comprising the Fischer-Tropsch product outof the reactor, the method comprising: (A) stopping the flow ofsynthesis gas into the reactor and the flow of effluent out of thereactor for a period of time of up to about 48 hours, or up to about 36hours, or up to about 24 hours, or up to about 12 hours, or up to about6 hours, or up to about 3 hours, or up to about 1 hour, or from about0.01 to about 48 hours, or from about 0.01 to about 36 hours, or fromabout 0.01 to about 24 hours, or from about 0.01 to about 12 hours, orfrom about 0.01 to about 6 hours, or from about 0.01 to about 3 hours,or from about 0.01 to about 1 hour, or from about 0.01 to about 0.5hour, or from about 0.1 to about 48 hours, or from about 0.1 to about 36hours, or from about 0.1 to about 24 hours, or from about 0.1 to about12 hours, or from about 0.1 to about 6 hours, or from about 0.1 to about3 hours, or from about 0.1 to about 2 hours, or from about 0.1 to about1 hour, or from about 0.1 to about 0.5 hour; (B) rejuvenating thecatalyst; and (C) restarting the flow of synthesis gas into the reactorand the flow of the effluent out of the reactor, the temperature of thesynthesis gas flowing into the reactor being at the desired reactiontemperature, or up to about 30° C., or up to about 20° C., or up toabout 10° C., or about 5° C., or about 2° C., or about 1° C. lower thanthe desired reaction temperature; the temperature within the reactorincreasing to the desired reaction temperature within a period of timeof up to about 2 hours, or up to about 1 hour, or up to about 0.5 hour,from the time of restarting the flow of synthesis gas into the reactor.In an embodiment, the reaction temperature is controlled with a heatexchange fluid flowing in a heat exchanger in thermal contact with thereactor, the temperature of the heat exchange fluid in the heatexchanger being lower than the desired reaction temperature.

The invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas conversion product and flowingeffluent comprising the synthesis gas conversion product out of thereactor, the method comprising: (A) stopping the flow of synthesis gasinto the reactor and the flow of effluent out of the reactor for aperiod of time of up to about 48 hours, or up to about 36 hours, or upto about 24 hours, or up to about 12 hours, or up to about 6 hours, orup to about 3 hours, or up to about 1 hour, or from about 0.01 to about48 hours, or from about 0.01 to about 36 hours, or from about 0.01 toabout 24 hours, or from about 0.01 to about 12 hours, or from about 0.01to about 6 hours, or from about 0.01 to about 3 hours, or from about0.01 to about 1 hour, or from about 0.01 to about 0.5 hour, or fromabout 0.1 to about 48 hours, or from about 0.1 to about 36 hours, orfrom about 0.1 to about 24 hours, or from about 0.1 to about 12 hours,or from about 0.1 to about 6 hours, or from about 0.1 to about 3 hours,or from about 0.1 to about 2 hours, or from about 0.1 to about 1 hour,or from about 0.1 to about 0.5 hour; (B) flowing hydrogen into thereactor in contact with the catalyst at a temperature of up to about400° C.; (C) flowing air into the reactor in contact with the catalystat a temperature in the range from about 70° C. to about 350° C., orabout 100° C. to about 350° C., or about 150° C. to about 350° C., orabout 200° C. to about 350° C., or about 250° C. to about 300° C., for aperiod of time in the range from about 1 to about 24 hours, or fromabout 1 to about 12 hours; (D) flowing hydrogen into the reactor incontact with the catalyst at a temperature of up to about 400° C. toregenerate the catalyst; and (E) restarting the flow of the synthesisgas into the reactor and the flow of the effluent out of the reactor.

This invention relates to a method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a microchannel reactor in contactwith a synthesis gas conversion catalyst at a desired reactiontemperature and pressure to produce a synthesis gas conversion productand flowing effluent comprising the synthesis gas conversion product outof the reactor, the method comprising: (A) stopping the flow of thesynthesis gas into the reactor and the flow of the effluent out of thereactor for up to 48 hours; and (B) restarting the flow of synthesis gasinto the reactor and the flow of effluent out of the reactor. In anembodiment, the flow of synthesis gas into the reactor and the flow ofeffluent out of the reactor is stopped for 0.01 to 48 hours. In anembodiment, prior to step (A) the pressure within the reactor is at apre-stoppage pressure, and during step (A) the pressure within thereactor is reduced to a level lower than the pre-stoppage pressure, andprior to step (B) the pressure within the reactor is increased to thepre-stoppage pressure. In an embodiment, subsequent to step (A) andprior to step (B) the reactor is purged by flowing hydrogen, natural gasor nitrogen gas in the reactor to purge the reactor. In an embodiment,the synthesis gas conversion process comprises a process for convertingsynthesis gas to methane, methanol or dimethyl ether. In an embodiment,the synthesis gas conversion process is a Fischer-Tropsch process. In anembodiment, the catalyst comprises cobalt. In an embodiment, thecatalyst comprises cobalt on a surface modified support. In anembodiment, the catalyst is in the form of a fixed bed of particulatesolids. In an embodiment, the synthesis gas conversion process is aFischer-Tropsch process and a tail gas is produced in the reactor, atleast part of the tail gas being combined with the synthesis gas to forma reactant mixture, the volumetric ratio of the synthesis gas to tailgas being in the range from 1:1 to 10:1. In an embodiment, the synthesisgas conversion process is a Fischer-Tropsch process and the synthesisgas comprising H₂ and CO, the mole ratio of H₂ to CO being in the rangefrom 1.4:1 to 2.1:1. In an embodiment, the reactor comprises one or morelayers of process microchannels and one or more layers of heat exchangechannels.

In any of the foregoing embodiments, the synthesis gas comprises CO andprior to stopping the flow of synthesis gas into the reactor theconversion of CO is at a desired conversion value, and after restartingthe flow of synthesis gas into the reactor the conversion of CO at thedesired conversion value is achieved within a time period of up to about3 hours, or up to about 2 hours, or up to about 1 hour, or up to about0.1 hour.

In any of the foregoing embodiment, the catalyst prior to step (A) maybe a wet catalyst.

In any of the foregoing embodiments, the reactor may comprise a fixedbed reactor, a fluidized bed reactor or a slurry phase reactor. Thereactor may comprise a conventional reactor. The reactor may comprise amicrochannel reactor. This invention may be especially valuable forreactor and process designs with small heat transfer temperaturedifferentials (e.g. average temperature differences between the reactiontemperature and the heat transfer fluid removing heat from the reactor),for example, less than about 10° C., or less than about 5° C., or lessthan about 2° C. These ranges for heat transfer temperaturedifferentials may offer advantages in process design (such as enablingcoolant wall temperatures to be above the water dew point at lowreaction temperatures), but may have a disadvantage in that processinterruptions may lead to catalyst deactivation unless the inventivemethods disclosed herein are used.

In any of the foregoing embodiments, the synthesis gas conversionprocess may comprise a process for converting synthesis gas to methane,methanol or dimethyl ether.

In any of the foregoing embodiments, the synthesis gas conversionprocess may comprise a Fischer-Tropsch process.

In any of the foregoing embodiments, the synthesis gas may comprise COand H₂ and the deactivation rate of the catalyst is less than a loss ofabout 0.2% CO conversion per day.

In any of the foregoing embodiments, the activity temperature delta forthe process after restarting the flow of synthesis gas into the reactormay be up to about 5° C., or up to about 3° C., or up to about 1° C.

In any of the foregoing embodiments, the relative activity ratio for aprocess utilizing this invention may be at least about 0.85, or at leastabout 0.92, or at least about 0.97, or at least about 1.00

In any of the foregoing embodiments, prior to the step of restarting theflow of synthesis gas into the reactor the temperature within thereactor may be below the desired reaction temperature, and thetemperature of the reactor may be increased to the desired reactiontemperature at a rate of up to about 5° C. per hour, or up to about 10°C. per hour, or up to about 15° C. per hour, or up to about 30° C. perhour, or up to about 60° C. per hour.

In any of the foregoing embodiments, the temperature in the reactor maybe at a pre-stoppage temperature prior to stopping the flow of synthesisgas into the reactor, and the temperature of the reactor during the stepof restarting the flow of synthesis gas into the reactor is at thepre-stoppage temperature.

In any of the foregoing embodiments, the temperature in the reactor maybe at a pre-stoppage temperature prior to stopping the flow of synthesisgas into the reactor, and the temperature of the reactor during the stepof restarting the flow of synthesis gas into the reactor is below thepre-stoppage temperature, for example, up to about 5° C. below thepre-stoppage temperature, or up to about 10° C. below the pre-stoppagetemperature, or up to about 15° C. below the pre-stoppage temperature,or up to about 20° C. below the pre-stoppage temperature.

In any of the foregoing embodiments, prior to stopping the flow ofsynthesis gas into the reactor, the temperature in the reactor may be ata desired operating temperature, and during the period of time betweenstopping the flow of synthesis gas into the reactor and restarting theflow of synthesis gas into the reactor the temperature in the reactormay be within about 20° C. of the desired operating temperature.

In any of the foregoing embodiments where a purge is employed, the purgemay be at a flow rate lower than or equal to the synthesis gas feed rateto the reactor prior to stopping the flow of synthesis gas into thereactor. In an embodiment, the total volume of purge gas may be lowerthan or equal to the total volume of synthesis gas between the locationsof stoppage of flow of synthesis gas into the reactor and the flow ofeffluent gas out of the reactor.

In any of the foregoing embodiments where a purge is employed, the purgemay be at a flow rate equal to or higher than the synthesis gas feedrate to the reactor prior to stopping the flow of synthesis gas into thereactor. In an embodiment, the total volume of purge gas may be higherthan the total volume of synthesis gas between the locations of stoppageof flow of synthesis gas into the reactor and the flow of effluent gasout of the reactor.

The synthesis gas conversion reactor (e.g., Fischer-Tropsch reactor) maycomprise a conventional reactor, or a microchannel reactor. Thesynthesis gas conversion reactor may comprise a fixed bed reactor, afluidized bed reactor or a slurry phase reactor. In each of theembodiments described above, microchannel reactors are highlyadvantageous and preferred due to enhanced heat transfer characteristicsprovided by the microchannel reactors. This is especially true for aFischer-Tropsch reaction which is highly exothermic, and themicrochannel reactor can provide a significantly higher level of heattransfer to control the Fischer-Tropsch reaction when compared to aconventional reactor.

BRIEF DESCRIPTION OF THE DRAWINGS

In the annexed drawings like parts and features have like references. Anumber of the drawings are schematic illustrations which may notnecessarily be drawn to scale.

FIG. 1 is a flow sheet illustrating a Fischer-Tropsch reaction processwhich comprises converting a fresh synthesis gas, which optionally maybe combined with a recycled tail gas, to a Fischer-Tropsch product. Thereaction process illustrated in FIG. 1 may be referred to as a reactiontrain.

FIG. 2 is a flow sheet illustrating a Fischer-Tropsch reaction processwhich comprises conducting the Fischer-Tropsch reaction in a pluralityof reaction trains (three reaction trains being shown in the drawing)connected to a common synthesis gas feedstream.

FIG. 3 is a schematic illustration of a microchannel reactor that can beused with the synthesis gas conversion processes referred to above.

FIGS. 4 and 5 are schematic illustrations of a microchannel reactor corethat can be used in the microchannel reactor illustrated in FIG. 3 .

FIG. 6 is a schematic illustration of a layer of process microchannelsthat can be used in the microchannel reactor core illustrated in FIGS. 4and 5 . The layer of process microchannels comprises a waveformpositioned in a space between planar plates. The process microchannelscomprise channels formed by the sides of the waveform and the planarplates. The space between the planar plates may have the dimensions of amicrochannel (i.e., a height of up to about 10 mm) and may also bereferred to as a microchannel or a process microchannel. A synthesis gasconversion catalyst (e.g., a Fischer-Tropsch catalyst) may be positionedin the process microchannels.

FIG. 7 is a schematic illustration of a layer of heat exchange channelsthat can be used in the microchannel reactor core illustrated in FIGS. 4and 5 . The heat exchange channels may be microchannels. As illustratedin FIGS. 5-7 , the heat exchange channels are oriented to provide for aflow of heat exchange fluid that is cross-current relative to the flowof fluid in the layer of process microchannels. However, the orientationof the heat exchange channels may be altered to provide for a flow ofheat exchange fluid that is co-current or counter-current relative tothe flow of fluid in the layer of process microchannels.

FIGS. 8-13 are schematic illustrations of catalysts or catalyst supportsthat may be used in the synthesis gas conversion reactor (e.g.,Fischer-Tropsch reactor). The catalyst illustrated in FIG. 8 is in theform of a fixed bed of particulate solids. The catalyst illustrated inFIG. 9 has a flow-by structure design. The catalyst illustrated in FIG.10 has a flow-through structure. FIGS. 11-13 are schematic illustrationsof fin assemblies that may be used for supporting the catalyst.

FIG. 14 is a schematic illustration of a synthesis gas conversionreactor (e.g., Fischer-Tropsch reactor) employing catalyst inserts inthe form of corrugated structures.

FIGS. 15-20 are charts showing the results of synthesis gas interruptiontests discussed below in Examples 1-6.

FIG. 21 is a flow sheet showing apparatus for conducting a synthesis gasconversion process which can be stopped and restarted.

DETAILED DESCRIPTION

All ranges and ratio limits disclosed in the specification and claimsmay be combined in any manner. It is to be understood that unlessspecifically stated otherwise, references to “a,” “an,” and/or “the” mayinclude one or more than one, and that reference to an item in thesingular may also include the item in the plural.

The phrase “and/or” should be understood to mean “either or both” of theelements so conjoined, i.e., elements that are conjunctively present insome cases and disjunctively present in other cases. Other elements mayoptionally be present other than the elements specifically identified bythe “and/or” clause, whether related or unrelated to those elementsspecifically identified unless clearly indicated to the contrary. Thus,as a non-limiting example, a reference to “A and/or B,” when used inconjunction with open-ended language such as “comprising” can refer, inone embodiment, to A without B (optionally including elements other thanB); in another embodiment, to B without A (optionally including elementsother than A); in yet another embodiment, to both A and B (optionallyincluding other elements); etc.

The word “or” should be understood to have the same meaning as “and/or”as defined above. For example, when separating items in a list, “or” or“and/or” shall be interpreted as being inclusive, i.e., the inclusion ofat least one, but also including more than one, of a number or list ofelements, and, optionally, additional unlisted items. Only terms clearlyindicated to the contrary, such as “only one of” or “exactly one of,” ormay refer to the inclusion of exactly one element of a number or list ofelements. In general, the term “or” as used herein shall only beinterpreted as indicating exclusive alternatives (i.e. “one or the otherbut not both”) when preceded by terms of exclusivity, such as “either,”“one of,” “only one of,” or “exactly one of.”

The phrase “at least one,” in reference to a list of one or moreelements, should be understood to mean at least one element selectedfrom any one or more of the elements in the list of elements, but notnecessarily including at least one of each and every elementspecifically listed within the list of elements and not excluding anycombinations of elements in the list of elements. This definition alsoallows that elements may optionally be present other than the elementsspecifically identified within the list of elements to which the phrase“at least one” refers, whether related or unrelated to those elementsspecifically identified. Thus, as a non-limiting example, “at least oneof A and B” (or, equivalently, “at least one of A or B,” or,equivalently “at least one of A and/or B”) can refer, in one embodiment,to at least one, optionally including more than one, A, with no Bpresent (and optionally including elements other than B); in anotherembodiment, to at least one, optionally including more than one, B, withno A present (and optionally including elements other than A); in yetanother embodiment, to at least one, optionally including more than one,A, and at least one, optionally including more than one, B (andoptionally including other elements); etc.

The transitional words or phrases, such as “comprising,” “including,”“carrying,” “having,” “containing,” “involving,” “holding,” and thelike, are to be understood to be open-ended, i.e., to mean including butnot limited to.

The term “microchannel” refers to a channel having at least one internaldimension of height or width of up to about 10 millimeters (mm), or upto about 5 mm, or up to about 2 mm. The microchannel may comprise atleast one inlet and at least one outlet wherein the at least one inletis distinct from the at least one outlet. The length of the microchannelmay be at least about two times the height or width, or at least aboutfive times the height or width, or at least about ten times the heightor width. The internal height or width of the microchannel may be in therange of up to about 10 mm, or from about 0.05 to about 10 mm, or fromabout 0.05 to about 8 mm, or from about 0.05 to about 7 mm, or fromabout 0.05 to about 5 mm, or from about 0.05 to about 3 mm, or fromabout 0.05 to about 2 mm, or from about 0.05 to about 1.5 mm, or fromabout 1 to about 10 mm, or from about 1 to about 8 mm, or from about 1to about 7 mm, or from about 1 to about 5 mm, or from about 1 to about 3mm, or from about 1 to about 2 mm, or from about 1 to about 1.5 mm. Theother internal dimension of height or width may be of any dimension, forexample, up to about 5 meters, or about 0.001 to about 5 meters, orabout 0.001 to about 3 meters, or about 0.001 to about 2 meters, orabout 0.001 to about 1 meter, or about 0.01 to about 0.5 meter, or about1 to about 10 mm, or about 1 to about 8 mm, or about 1 to about 7 mm, orabout 1 to about 5 mm, or about 1 to about 3 mm, or about 1 to about 2mm. The length of the microchannel may be of any dimension, for example,up to about 5 meters, or from about 0.1 to about 5 meters, or from about0.1 to about 3 meters, or from about 0.1 to about 2 meters, or fromabout 0.1 to about 1.5 meters, or from about 0.1 to about 1 meter, orabout 0.15 to about 5 meters, or from about 0.15 to about 3 meters, orfrom about 0.15 to about 2.5 meters, or from about 0.15 to about 2meters, or from about 0.15 to about 1.5 meters, or from about 0.15 toabout 1 meter. The length may be in the range from about 0.1 to about0.8 meter, or from about 0.1 to about 0.6 meter, or from about 0.1 toabout 0.5 meter, or from about 0.1 to about 0.3 meter. The microchannelmay have a cross section having any shape, for example, a square,rectangle, circle, triangle, semi-circle, trapezoid, etc. The shapeand/or size of the cross section of the microchannel may vary over itslength. For example, the height or width may taper from a relativelylarge dimension (e.g., about 10 mm) to a relatively small dimension(e.g., about 1 mm), or vice versa, over the length of the microchannel.

The term “layer of microchannels” refers to a plurality of microchannelsaligned parallel to each other and positioned in a plane. The layer ofmicrochannels may be in the form of a layer of process microchannels ora layer of heat exchange microchannels. The layer of microchannels maycomprise a waveform positioned between planar plates. The microchannelsmay comprise channels formed by the sidewalls of the waveform and planarplates. The space between the planar plates may also have the dimensionsof a microchannel (i.e., a height of up to about 10 mm) and thereforemay be referred to as a microchannel. A catalyst may be positioned inthe process microchannels.

The term “microchannel reactor” refers to an apparatus comprising one ormore layers of process microchannels wherein a process is conducted. Theprocess may be a synthesis gas conversion process, for example, aFischer-Tropsch (FT) reaction process. The microchannel reactor mayfurther comprise a heat exchanger, for example, one or more layers ofheat exchange channels adjacent to and/or in thermal contact with theone or more layers of process microchannels. The heat exchange channelsmay provide cooling for the fluids in the process microchannels.

The heat exchange channels may be microchannels. The layers of processmicrochannels and heat exchange channels may be stacked one above theother or positioned side-by-side to form a microchannel reactor core;see, FIGS. 4 and 5 . The microchannel reactor core may have the form ofa three-dimensional block which has six faces that may be squares orrectangles. The microchannel reactor core may have the samecross-section along a length. The microchannel reactor core may have theshape of a rectangular or cubic block. The microchannel reactor core maybe in the form of a rectangular prism or a cube. The microchannelreactor core may have a length, width and height of any dimension, forexample, a length in the range from about 0.1 to about 5 meters, or fromabout 0.1 to about 3 meters, or from about 0.1 to about 2 meters, orfrom about 0.1 to about 1.5 meters, or from about 0.1 to about 1 meter,or from about 0.1 to about 0.75 meter, or from about 0.1 to about 0.5meter, or from about 0.1 to about 0.25 meter, or from about 0.1 to about5 meter. The microchannel core may have a width in the range from about0.1 to about 5 meters, or from about 0.1 to about 3 meters, or fromabout 0.1 to about 2 meters, or from about 0.1 to about 1.5 meters, orfrom about 0.1 to about 1 meter, or from about 0.1 to about 0.75 meter,or from about 0.1 to about 0.5 meter, or from about 0.1 to about 0.25meter, or from about 0.1 to about 5 meter. The microchannel reactor coremay have a height in the range from about 0.1 to about 5 meters, or fromabout 0.1 to about 3 meters, or from about 0.1 to about 2 meters, orfrom about 0.1 to about 1.5 meters, or from about 0.1 to about 1 meter,or from about 0.1 to about 0.75 meter, or from about 0.1 to about 0.5meter, or from about 0.1 to about 0.25 meter, or from about 0.1 to about5 meter. The microchannel reactor may include a header for providing forthe flow of fluid into the one or more layers of process microchannels,and a footer providing for the flow of fluid out of the one or morelayers of process microchannels. The microchannel reactor may include aheader for providing for the flow of heat exchange fluid into the heatexchange channels, and a footer providing for the flow of heat exchangefluid out of the one or more layers of heat exchange channels. Themicrochannel reactor may be sufficiently small and compact so as to bereadily transportable. As such, the reactor along with the otherequipment used in the synthesis gas conversion process (e.g., theFischer-Tropsch process) may be readily transported to remote locations,such as military bases, shale oil locations, coal mines, and the like,where the source of synthesis gas may be located. These reactors may beused on ships, oil drilling platforms, and the like.

The term “process microchannel” refers to a microchannel wherein aprocess is conducted. The process may be a synthesis gas conversionprocess, for example, a Fischer-Tropsch (FT) reaction process.

The term “conventional reactor” refers to a reactor that is not amicrochannel reactor.

The term “reaction train” is used herein to refer to an apparatus with aheat exchange fluid (e.g., a coolant) at a specific controlledtemperature, including one or more reactors and feed lines for flowingreactants into the one or more reactors and product or effluent linesfor flowing product and/or effluent out of the one or more reactors. Thereaction train may be used for conducting a synthesis gas conversionprocess, for example, a Fischer-Tropsch reaction process.

The term “plant” is used herein to refer to a facility employing one ormore reaction trains. When two or more reaction trains are employed, thereaction trains may be connected to a common source of reactants, e.g.,synthesis gas.

The term “volume” with respect to volume within a process microchannelmay include all volume in the process microchannel a process fluid mayflow through or flow by. This volume may include volume within surfacefeatures that may be positioned in the process microchannel and adaptedfor the flow of fluid in a flow-through manner or in a flow-by manner.

The term “adjacent” when referring to the position of one channelrelative to the position of another channel may mean directly adjacentsuch that a wall or walls separate the two channels. In one embodiment,the two channels may have a common wall. The common wall may vary inthickness. However, “adjacent” channels may not be separated by anintervening channel that may interfere with heat transfer between thechannels. One channel may be adjacent to another channel over only partof the dimension of the another channel. For example, a processmicrochannel may be longer than and extend beyond one or more adjacentheat exchange channels.

The term “thermal contact” refers to two bodies, for example, twochannels, that may or may not be in physical contact with each other oradjacent to each other but still exchange heat with each other. One bodyin thermal contact with another body may heat or cool the other body. Alayer of process microchannels may be in thermal contact with a layer ofheat exchange channels.

The term “fluid” refers to a gas, a liquid, a mixture of a gas and aliquid, or a gas or a liquid containing dispersed solids, liquiddroplets and/or gaseous bubbles. The droplets and/or bubbles may beirregularly or regularly shaped and may be of similar or differentsizes.

The term “effluent” is used herein to refer to a process fluid (not aheat exchange fluid) flowing out of a reactor. The effluent may compriseunreacted reactants, diluents, inerts, product, or a mixture of two ormore thereof.

The terms “gas” and “vapor” may have the same meaning and are sometimesused interchangeably.

The term “residence time” or “average residence time” refers to theinternal volume of a space within a channel occupied by a fluid flowingin the space divided by the average volumetric flow rate for the fluidflowing in the space at the temperature and pressure being used.

The term “activity temperature delta” refers to an increase in thereaction temperature required to maintain the conversion rate of CO in asynthesis gas conversion process (e.g., Fischer-Tropsch reactionprocess) at a constant level. The reaction temperature may be thetemperature of the process fluid (e.g., reactants, product, etc.) as itleaves the catalyst. The activity temperature delta can be calculated asa change in productivity of the catalyst (e.g., volume of CO convertedper unit volume of catalyst per hour) at conditions (flows, feedcomposition and reactor temperature) that existed prior to the upset,i.e., prior to the stop of the flow of synthesis gas into the reactor.The amount of temperature increase needed to compensate for thisactivity loss would be dependent on the apparent activation energy ofthe catalyst in the reactor under consideration. For a Fischer-Tropschcatalyst operating under kinetically limited regime with an apparentactivation energy of 100 kJ/mol, a 2% loss in activity may result in anequivalent 2% reduction in reaction rates and may necessitate atemperature increase of approximately 1.0° C. to restore thepre-stoppage catalyst productivity under identical operating conditions.In case of change in operating conditions, this may be assessed by useof appropriate rate expression that includes dependence on the partialpressure of the reaction components. An example rate expression for aFischer-Tropsch catalyst would be “Intrinsic kinetics of theFischer-Tropsch synthesis on a cobalt catalyst, Ian C. Yates and CharlesN. Satterfield, Energy Fuels, 1991, 5 (1), pp 168-173.

The term “relative activity ratio” refers to a ratio of the activityfactor post the process restart to that before the process stoppage. Theactivity factor can be calculated as a ratio of the estimated reactionrate from process data and the predicted reaction rate using anestablished model that defines the apparent kinetic rate. For example,in the case of Fischer-Tropsch synthesis, the activity factor can becalculated as the ratio of reaction rate estimated as the mols of COconverted per unit volume of catalyst per hour (using synthesis gas flowrate and composition and measured CO conversion) and the model predictedreaction rate (in moles of CO converted per unit volume of the catalystper hour) calculated by the kinetic rate equation from “IntrinsicKinetics of the Fischer-Tropsch Synthesis on a Cobalt Catalyst,” Ian C.Yates and Charles N. Satterfield, Energy Fuels, 1991, 5 (1), pp 168-173.

The terms “upstream” and “downstream” refer to positions within achannel (e.g., a process microchannel) or in a process flow sheet thatis relative to the direction of flow of a fluid in the channel orprocess flow sheet. For example, a position within a channel or processflow sheet not yet reached by a portion of a fluid stream flowing towardthat position would be downstream of that portion of the fluid stream. Aposition within the channel or process flow sheet already passed by aportion of a fluid stream flowing away from that position would beupstream of that portion of the fluid stream. The terms “upstream” and“downstream” do not necessarily refer to a vertical position since thechannels used herein may be oriented horizontally, vertically or at aninclined angle.

The term “plate” refers to a planar or substantially planar sheet ofmaterial. The plate may be referred to as a shim. The thickness of theplate may be the smallest dimension of the plate and may be up to about4 mm, or in the range from about 0.05 to about 4 mm, or in the rangefrom about 0.05 to about 2 mm, or in the range of about 0.05 to about 1mm, or in the range from about 0.05 to about 0.5 mm. The plate may haveany length and width.

The term “surface feature” refers to a depression in a channel walland/or a projection from a channel wall that disrupts flow within thechannel. The surface features may be in the form of circles, spheres,frustrums, oblongs, squares, rectangles, angled rectangles, checks,chevrons, vanes, airfoils, wavy shapes, and the like, and combinationsof two or more thereof. The surface features may contain subfeatureswhere the major walls of the surface features further contain smallersurface features that may take the form of notches, waves, indents,holes, burrs, checks, scallops, and the like. The surface features mayhave a depth, a width, and for non-circular surface features a length.The surface features may be formed on or in one or more of the interiorwalls of the process microchannels used in accordance with the disclosedprocess. The surface features may be referred to as passive surfacefeatures or passive mixing features. The surface features may be used todisrupt flow (for example, disrupt laminar flow streamlines) and createadvective flow at an angle to the bulk flow direction.

The term “heat exchange channel” refers to a channel having a heatexchange fluid in it that provides heat and/or absorbs heat. The heatexchange channel may absorb heat from or provide heat to an adjacentchannel (e.g., process microchannel) and/or one or more channels inthermal contact with the heat exchange channel. The heat exchangechannel may absorb heat from or provide heat to channels that areadjacent to each other but not adjacent to the heat exchange channel. Inone embodiment, one, two, three or more channels may be adjacent to eachother and positioned between two heat exchange channels. The heatexchange channel may be a cooling channel.

The term “heat transfer wall” refers to a common wall between a layer ofprocess microchannels and an adjacent layer of heat exchange channelswhere heat transfers from one channel layer to the other through thecommon wall.

The term “heat exchange fluid” refers to a fluid that may give off heatto and/or absorb heat from the surrounding walls of a heat exchangechannel via conduction, convection or phase change (e.g., giving offheat while transforming from vapor to liquid or absorbing heat whiletransforming from liquid to vapor). The heat exchange fluid may be in orflow in a heat exchange channel or a layer of heat exchange channels.The term “heat exchange fluid” may be used interchangeably with the word“coolant.”

The term “waveform” refers to a contiguous piece of sheet material thatis transformed from a planar object to a three-dimensional object. Thewaveform may be used to form one or more microchannels. The waveform maycomprise a right angled corrugated sheet which may be sandwiched betweenopposed planar plates. The right angled corrugated sheet may haverounded edges. The waveform may have a sinusoidal shape; see, FIGS. 5and 6 . One or more microchannels in a microchannel layer may be definedon three sides by the waveform and on the fourth side by one of theplanar plates. The waveform may be made of any of the materialsdisclosed herein as being useful for making the microchannel reactor.These may include copper, aluminum, stainless steel, and the like. Thesemay also include overlay, inlaid, or edge clad materials which combinetwo alloys (for example by cold rolling) to produce a clad material withproperties which are different than either alloy alone. The thermalconductivity of the waveform may be about 1 W/m-K or higher.

The term “bulk flow direction” refers to the vector through which fluidmay travel in an open path in a channel.

The term “bulk flow region” refers to open areas within a microchannel.A contiguous bulk flow region may allow rapid fluid flow through amicrochannel without significant pressure drops. In one embodiment, theflow in the bulk flow region may be laminar. A bulk flow region maycomprise at least about 5% of the internal volume and/or cross-sectionalarea of a microchannel, or from about 5% to about 100%, or from about 5%to about 99%, or from about 5% to about 95%, or from about 5% to about90%, or from about 30% to about 80% of the internal volume and/orcross-sectional area of the microchannel.

The terms “open channel” or “flow-by channel” or “open path” refers to achannel (e.g., a microchannel) with a gap with a height of at leastabout 0.01 mm that extends through the channel such that fluid may flowthrough the channel without encountering a barrier to flow. The gap mayhave a height of up to about 10 mm, or up to about 5 mm, or up to about2 mm, or up to about 1 mm, or up to about 0.5 mm.

The term “cross-sectional area” of a channel (e.g., processmicrochannel) refers to an area measured perpendicular to the directionof the bulk flow of fluid in the channel and may include all areaswithin the channel including any surface features that may be present,but does not include the channel walls. For channels that curve alongtheir length, the cross-sectional area may be measured perpendicular tothe direction of bulk flow at a selected point along a line thatparallels the length and is at the center (by area) of the channel.Dimensions of height and width may be measured from one channel wall tothe opposite channel wall. These dimensions may not be changed byapplication of a coating to the surface of the wall. These dimensionsmay be average values that account for variations caused by surfacefeatures, surface roughness, and the like.

The term “open cross-sectional area” of a channel (e.g., processmicrochannel) refers to an area open for bulk fluid flow in a channelmeasured perpendicular to the direction of the bulk flow of fluid flowin the channel. The open cross-sectional area may not include internalobstructions such as surface features and the like which may be present.

The term “superficial velocity” for the velocity of a fluid flowing in achannel refers to the velocity resulting from dividing the volumetricflow rate of the fluid at the inlet temperature and pressure of thechannel by the cross-sectional area of the channel.

The term “free stream velocity” refers to the velocity of a streamflowing in a channel at a sufficient distance from the sidewall of thechannel such that the velocity is at a maximum value. The velocity of astream flowing in a channel is zero at the sidewall if a no slipboundary condition is applicable, but increases as the distance from thesidewall increases until a constant value is achieved. This constantvalue is the “free stream velocity.”

The term “bottled” is used herein to refer to a process shut downprocedure wherein the flow of reactants into a reactor and the flow ofeffluent out of the reactor are stopped. The term “isolated” may be usedin place of “bottled.” The net effect is the isolation of a processfluid (e.g. synthesis gas) in the reactor in contact with the catalyst.

The term “process fluid” is used herein to refer to reactants, productand any diluent or other fluid that may flow in a process microchannel.

The term “reaction zone” refers to the space within a microchannelwherein a chemical reaction occurs or wherein a chemical conversion ofat least one species occurs. The reaction zone may contain one or morecatalysts.

The term “contact time” refers to the volume of a reaction zone within amicrochannel divided by the volumetric feed flow rate of the reactantsat a temperature of 0° C. and a pressure of one atmosphere.

The term “fresh synthesis gas” refers to synthesis gas that flows into amicrochannel reactor and is used as a reactant in a synthesis gasconversion process (e.g., a Fischer-Tropsch reaction). Synthesis gascomprises a mixture of CO and H₂. Synthesis gas may be referred to assyngas. Fresh synthesis gas is not a FT tail gas.

The term “tail gas” refers to a gaseous product produced during asynthesis gas conversion process. For example, the tail gas may be a “FTtail gas” which is a tail gas produced during a Fisher-Tropsch reaction.The tail gas may contain CO and H₂. The tail gas may be combined withfresh synthesis gas to form a reactant mixture.

The term “reactant mixture” refers to a mixture of fresh synthesis gasand tail gas recycled from a synthesis gas conversion process (e.g., aFischer-Tropsch reaction).

The term “conversion of CO” refers to the CO mole change between thefresh synthesis gas in a reactant feed stream and the effluent gasleaving the reactor, divided by the moles of CO in the fresh synthesisgas in the reactant feed stream.

The term “one-pass conversion of CO” refers to the conversion of CO fromthe overall reactant mixtures (i.e., fresh synthesis gas plus recycledtail gas or recycled tail gas components) after one pass through themicrochannel reactor.

The term “selectivity to methane” refers to the moles of methane in aproduct minus the moles of methane in a reactant, divided by the molesof CO in the reactant that are consumed in the reaction.

The term “yield” refers to the number of moles of product exiting areactor divided by the number of moles of a reactant entering thereactor.

The term “cycle” refers to a single pass of the reactants through areactor.

The term “wet catalyst” refers to a catalyst (e.g., a catalyst bed ofparticulate solids) which has produced liquid product, for exampleFischer-Tropsch liquid hydrocarbon product, and has a liquid film on itssurface and/or in its pores that may increase the diffusional resistancefor gas phase reactants to reach active catalyst sites, slowing down theapparent reaction rate relative to that for a catalyst not containingsuch a film. Use of the term “wax” to describe a Fischer-Tropsch productimplies that the material is a liquid while in the reactor at thereaction conditions of temperature and pressure.

The term “graded catalyst” refers to a catalyst with one or moregradients of catalytic activity. The graded catalyst may have a varyingconcentration or surface area of a catalytically active metal. Thegraded catalyst may have a varying turnover rate of catalytically activesites. The graded catalyst may have physical properties and/or a formthat varies as a function of distance. For example, the graded catalystmay have an active metal concentration that is relatively low at theentrance to a process microchannel and increases to a higherconcentration near the exit of the process microchannel, or vice versa;or a lower concentration of catalytically active metal nearer the center(i.e., midpoint) of a process microchannel and a higher concentrationnearer a process microchannel wall, or vice versa, etc. The thermalconductivity of a graded catalyst may vary from one location to anotherwithin a process microchannel. The surface area of a graded catalyst maybe varied by varying size of catalytically active metal sites on aconstant surface area support, or by varying the surface area of thesupport such as by varying support type or particle size. A gradedcatalyst may have a porous support where the surface area to volumeratio of the support is higher or lower in different parts of theprocess microchannel followed by the application of the same catalystcoating everywhere. A combination of two or more of the precedingembodiments may be used. The graded catalyst may have a single catalyticcomponent or multiple catalytic components (for example, a bimetallic ortrimetallic catalyst). The graded catalyst may change its propertiesand/or composition gradually as a function of distance from one locationto another within a process microchannel. The graded catalyst maycomprise rimmed particles that have “eggshell” distributions ofcatalytically active metal within each particle. The graded catalyst maybe graded in the axial direction along the length of a processmicrochannel or in the lateral direction. The graded catalyst may havedifferent catalyst compositions, different loadings and/or numbers ofactive catalytic sites that may vary from one position to anotherposition within a process microchannel. The number of catalyticallyactive sites may be changed by altering the porosity of the catalyststructure. This may be accomplished using a washcoating process thatdeposits varying amounts of catalytic material. An example may be theuse of different porous catalyst thicknesses along the processmicrochannel length, whereby a thicker porous structure may be leftwhere more activity is required. A change in porosity for a fixed orvariable porous catalyst thickness may also be used. A first pore sizemay be used adjacent to an open area or gap for flow and at least onesecond pore size may be used adjacent to the process microchannel wall.

The term “Fischer-Tropsch” or “FT” refers to a chemical reactionrepresented by the equation:nCO+2nH₂→(CH₂)_(n) +nH₂OThis reaction is an exothermic reaction. n may be any number, forexample from 1 to about 1000, or from about 2 to about 200, or fromabout 5 to about 150.

The term “synthesis gas conversion product” refers to a product formedby a synthesis gas conversion process. The product may comprise methane,methanol and/or dimethyl ether. The synthesis gas conversion product maybe a Fischer-Tropsch product.

The term “Fischer-Tropsch product” or “FT product” refers to ahydrocarbon product made by a Fischer-Tropsch process. The FT productmay comprise a liquid, solid, gas, or mixture of two or more thereof.The FT solid product may be in the form of a wax, i.e., aFischer-Tropsch wax or FT wax. The FT wax may be melted or hydrocrackedto form a FT liquid. The FT liquid product may have a boiling point ator above about 30° C. at atmospheric pressure. The FT gas may bereferred to as a FT tail gas. The “FT tail gas” may have a boiling pointbelow about 30° C. at atmospheric pressure. The FT tail gas may containH₂ and CO.

The term “dry” catalyst refers to a fresh catalyst or a regeneratedcatalyst that has not been exposed to a synthesis gas conversion processand does not have a liquid and/or solid synthesis gas conversion productadhered to its surface or in its pores.

The term “rejuvenate a catalyst” refers to a process for removingprocess liquid and/or solids (e.g., FT liquid and/or FT wax) from thesurface of a catalyst and/or the pores of the catalyst. The catalyst maybe rejuvenated using hydrogen or a solvent. The solvent may be a liquidhydrocarbon solvent.

The term “regenerate a catalyst” refers to a process wherein processliquid and/or solids (e.g., FT liquid and/or FT wax) are removed fromthe surface of a catalyst and/or the pores of the catalyst, and then thecatalyst is subjected to oxidation and reduction. Alternatively, thecatalyst may be regenerated using hydrogen.

The term “desulfurized natural gas” refers to a natural gas that hasbeen processed in a desulfurization step. The desulfurized natural gasmay have a sulfur content of up to about 100 parts per billion by volume(ppbv), or up to about 50 ppbv, or up to about 30 ppbv, or up to about10 ppbv, or up to about 5 ppbv.

The term “chain growth” refers to the growth in a molecule resultingfrom a reaction in which the molecule grows with the addition of newmolecular structures (e.g., the addition of methylene groups to ahydrocarbon chain in a Fischer-Tropsch synthesis).

The term “aliphatic hydrocarbon” refers to aliphatic compounds, such asalkanes, alkenes, alkynes, and the like.

The term “higher molecular weight aliphatic hydrocarbon” refers to analiphatic hydrocarbon having 2 or more carbon atoms, or 3 or more carbonatoms, or 4 or more carbon atoms, or 5 or more carbon atoms, or 6 ormore carbon atoms.

The higher molecular weight aliphatic hydrocarbons may have up to about200 carbon atoms or higher, or up to about 150 carbon atoms, or up toabout 100 carbon atoms, or up to about 90 carbon atoms, or up to about80 carbon atoms, or up to about 70 carbon atoms, or up to about 60carbon atoms, or up to about 50 carbon atoms, or up to about 40 carbonatoms, or up to about 30 carbon atoms. Examples may include ethane,propane, butane, pentane, hexane, octane, decane, dodecane, and thelike.

The term “Co loading” refers to the weight of Co in a catalyst dividedby the total weight of the catalyst, that is, the total weight of the Coplus any co-catalyst or promoter as well as any support. If the catalystis in the form of particulate solids which include a support material(e.g., Al₂O₃), the weight of the support material is included in thecalculation. If the catalyst is supported on an engineered supportstructure, such as a foam, felt, wad or fin, the weight of suchengineered support structure is not included in the calculation.Similarly, if the catalyst is adhered to a channel wall, the weight ofthe channel wall is not to be included in the calculation.

The term “mm” may refer to millimeter. The term “nm” may refer tonanometer. The term “ms” may refer to millisecond. The term “ρs” mayrefer to microsecond. The term “μm” may refer to micron or micrometer.The terms “micron” and “micrometer” have the same meaning and may beused interchangeably.

Unless otherwise indicated, all pressures are expressed in terms ofabsolute pressure.

The synthesis gas conversion process (e.g., Fischer-Tropsch process)employed herein may use fresh synthesis gas as a synthesis gasconversion reactant (e.g., Fischer-Tropsch reactant). Optionally, a FTtail gas may be used in combination with the fresh synthesis gas to forma reactant mixture. The fresh synthesis gas may be produced using steamreforming (e.g., a steam methane reforming (SMR) reaction where methaneis reacted with steam in the presence of a steam methane reforming (SMR)catalyst); partial oxidation; autothermal reforming; carbon dioxidereforming; or a combination of two or more thereof.

The fresh synthesis gas may be produced by gasifying a carbonaceousmaterial at an elevated temperature, for example, about 700° C. orhigher. The carbonaceous material may comprise any carbon-containingmaterial that can be gasified to produce synthesis gas. The carbonaceousmaterial may comprise natural gas, biomass (e.g., plant or animalmatter, biodegradable waste, and the like), a food resource (e.g., ascorn, soybean, and the like), and/or a non-food resource such as coal(e.g., low grade coal, high grade coal, clean coal, and the like), oil(e.g., crude oil, heavy oil, tar sand oil, shale oil, and the like),solid waste (e.g., municipal solid waste, hazardous waste), refusederived fuel (RDF), tires, petroleum coke, trash, garbage, biogas,sewage sludge, animal waste, agricultural waste (e.g., corn stover,switch grass, grass clippings), construction demolition materials,plastic materials (e.g., plastic waste), cotton gin waste, landfill gas,a mixture of two or more thereof, and the like.

The fresh synthesis gas may contain water and/or particulate solidswhich may be separated from the fresh synthesis gas before flowing thefresh synthesis gas into a synthesis gas conversion reactor (e.g., aFischer-Tropsch reactor). If the synthesis gas conversion reactor is amicrochannel reactor, the presence of such solids and/or water in thefresh synthesis gas may be detrimental to the operation of themicrochannel reactor due to the fact that the passages in themicrochannel reactor are very small (e.g., a process microchannel has aninternal height or width of up to only about 10 mm). The removal of (orsignificant reduction in the concentration) of such water and/orparticulate solids may avoid or reduce the likelihood of a reduction inthe reaction rate of the synthesis gas conversion process (e.g.,Fischer-Tropsch process), as well as avoid clogging in the processmicrochannels. It can also avoid premature catalyst deactivation. Thisis significant due to the fact that the process microchannels of amicrochannel reactor use less catalyst than a conventional (i.e.,non-microchannel) reactor.

The fresh synthesis gas may comprise H₂ and CO with the molar ratio ofH₂ to CO in the range from about 1.9:1 to about 2.1:1, or from about1.95:1 to about 2.05:1, or from about 1.98:1 to about 2.02:1.

The fresh synthesis gas may optionally be combined with a recycled tailgas (e.g., a recycled FT tail gas), which also contains H₂ and CO, toform a reactant mixture. The tail gas may comprise H₂ and CO with amolar ratio of H₂ to CO in the range from about 0.5:1 to about 2:1, orfrom about 0.6:1 to about 1.8:1, or from about 0.7:1 to about 1.2:1.

The reactant mixture may comprise H₂ and CO with a molar ratio of H₂ toCO that may be in the range from about 1.4:1 to about 2.1:1, or fromabout 1.5:1 to about 2:1:1, or from about 1.6:1 to about 2:1, or fromabout 1.7:1 to about 1.9:1.

When the recycled tail gas is used, the volumetric ratio of freshsynthesis gas to recycled tail gas used to form the reactant mixture maybe in the range from about 1:1 to about 10:1, or from about 1:1 to about8:1, or from about 1:1 to about 6:1, or from about 1:1 to about 4:1, orfrom about 3:2 to about 7:3, or about 2:1.

The reactor may be a conventional reactor or a microchannel reactor. Inan embodiment, the reactor may be characterized by a heat transfersurface (or heat transfer wall) for removing heat of reaction from thereactor (or process microchannel layer) wherein the ratio of the surfacearea of the heat transfer surface to the volume of the catalyst in thereactor is at least about 300 square meters (m²) of heat transfersurface per cubic meter (m³) of catalyst, or from about 300 to about5000 m²/m³; or from about 1000 to about 3000 m²/m³.

In the following discussion relating to the embodiments illustrated inFIGS. 1 and 2 , the synthesis gas conversion process will be discussedin terms of being a Fischer-Tropsch reaction process. However,alternatively the synthesis gas conversion process illustrated in FIGS.1 and 2 may be a process for converting synthesis gas to methane,methanol or dimethyl ether, and the discussion provided below isapplicable to such processes.

Referring to FIG. 1 , the Fischer-Tropsch process 100 employs the use ofFischer-Tropsch reactor 110. The Fischer-Tropsch reactor 110 may be aconventional reactor or a microchannel reactor, although a microchannelreactor is preferred. The Fischer-Tropsch reactor may be a fixed bedreactor, a fluidized bed reactor, or a slurry phase reactor. When amicrochannel reactor is used it may be referred to as a Fischer-Tropschor FT microchannel reactor. In operation, fresh synthesis gas 120 flowsinto the Fischer-Tropsch reactor 110. Optionally, the fresh synthesisgas may be combined with recycled tail gas 130 to form reactant mixture140 which flows into the Fischer-Tropsch reactor 110. The freshsynthesis gas may be combined with the recycled tail gas upstream of theFischer-Tropsch reactor 110, as shown in FIG. 1 , or in theFischer-Tropsch reactor 110.

In an embodiment, referring to FIG. 1 , the flow of synthesis gas intothe reactor 110 may be stopped, and the Fischer-Tropsch process may berestarted by a method comprising: restoring the pressure within thereactor 110 at the reaction pressure, maintaining the temperature withinthe reactor at the desired reaction temperature, and restarting the flowof synthesis gas into the reactor.

In an embodiment, the flow of synthesis gas into the reactor 110 isstopped, and the Fischer-Tropsch process is restarted by a methodcomprising: restoring the pressure within the reactor at the reactorreaction pressure, and restarting the flow of synthesis gas into thereactor. The reactor temperature may then ramped up to the operatingtemperature that was employed prior to the stop within a period of up toabout 36 hours, or up to about 24 hours, or up to about 12 hours, or upto about 6 hours, or up to about 3 hours, or from about 0.01 to about 48hours, or from about 0.01 to about 36 hours, or from about 0.01 to about24 hours, or from about 0.01 to about 12 hours, or from about 0.01 toabout 6 hours, or from about 0.01 to about 3 hours, or from about 0.01to about 1 hour, or from about 0.01 to about 0.5 hour, or from about 0.1to about 48 hours, or from about 0.1 to about 36 hours, or from about0.1 to about 24 hours, or from about 0.1 to about 12 hours, or fromabout 0.1 to about 6 hours, or from about 0.1 to about 3 hours, or fromabout 0.1 to about 3 hours, or from about 0.1 to about 1 hour, or fromabout 0.1 to about 0.5 hour by heating at a rate of up to about 5° C.per hour, or up to about 10° C. per hour, or up to about 15° C. perhour, or up to about 30° C. per hour, or up to about 60° C. per hour.The reaction temperature may be controlled with a coolant flowing in aheat exchanger in thermal contact with the reactor.

In an embodiment, the flow of synthesis gas into the reactor 110 isstopped, and the Fischer-Tropsch process may be restarted by a methodcomprising: flowing hydrogen into the reactor to purge the reactor ofreactants and effluent, holding the reactor in a hydrogen environmentfor a period of time of up to about 48 hours, or up to about 36 hours,or up to about 24 hours, or up to about 12 hours, or up to about 6hours, or up to about 3 hours, or up to about 1 hour, or from about 0.01to about 48 hours, or from about 0.01 to about 36 hours, or from about0.01 to about 24 hours, or from about 0.01 to about 12 hours, or fromabout 0.01 to about 6 hours, or from about 0.01 to about 3 hours, orfrom about 0.01 to about 1 hour, or from about 0.01 to about 0.5 hour,or from about 0.1 to about 48 hours, or from about 0.1 to about 36hours, or from about 0.1 to about 24 hours, or from about 0.1 to about12 hours, or from about 0.1 to about 6 hours, or from about 0.1 to about3 hours, or from about 0.1 to about 2 hours, or from about 0.1 to about1 hour, or from about 0.1 to about 0.5 hour; and restarting the flow ofsynthesis gas into the reactor. The purge may be at a flow rate lowerthan or equal to or greater than the synthesis gas feed rate to thereactor prior to stoppage. The reaction temperature may be in the rangefrom about 150° C. to about 300° C., and during the step of flowinghydrogen into the reactor the temperature within the reactor mayincrease to a temperature above the reaction temperature, for example,up to about 350° C., or up to about 400° C. The reaction temperature maybe controlled with a coolant flowing in a heat exchanger in thermalcontact with the reactor.

In an embodiment, the flow of synthesis gas into the reactor 110 isstopped, and the Fischer-Tropsch process may be restarted by a methodcomprising: flowing desulfurized natural gas into the reactor to purgethe reactor of reactants and effluent, holding the reactor in a naturalgas environment at a temperature in the range of about 150 to about 300°C. for a period of time of up to about 48 hours, or up to about 36hours, or up to about 24 hours, or up to about 12 hours, or up to about6 hours, or up to about 3 hours, or up to about 1 hour, or from about0.01 to about 48 hours, or from about 0.01 to about 36 hours, or fromabout 0.01 to about 24 hours, or from about 0.01 to about 12 hours, orfrom about 0.01 to about 6 hours, or from about 0.01 to about 3 hours,or from about 0.01 to about 1 hour, or from about 0.01 to about 0.5hour, or from about 0.1 to about 48 hours, or from about 0.1 to about 36hours, or from about 0.1 to about 24 hours, or from about 0.1 to about12 hours, or from about 0.1 to about 6 hours, or from about 0.1 to about3 hours, or from about 0.1 to about 2 hours, or from about 0.1 to about1 hour, or from about 0.1 to about 0.5 hour; and then restarting theflow of synthesis gas into the reactor and the flow of effluent out ofthe reactor. The purge may be at a flow rate lower than or equal to orgreater than the synthesis gas feed rate to the reactor prior tostoppage.

In an embodiment, the flow of synthesis gas into the reactor 110 isstopped, and the Fischer-Tropsch process may be restarted by a methodcomprising: maintaining the reactor at the pre-stop operatingtemperature for a period of up to about 12 hours, or up to about 6hours, or up to about 3 hours, or up to about 1 hour, then flowinghydrogen, nitrogen, or desulfurized natural gas into the reactor topurge the reactor of reactants and effluent, holding the reactor in apurge gas environment for a period of time of up to about 48 hours, orup to about 36 hours, or up to about 24 hours, or up to about 12 hours,or up to about 6 hours, or up to about 3 hours, or up to about 1 hour,or from about 0.01 to about 48 hours, or from about 0.01 to about 36hours, or from about 0.01 to about 24 hours, or from about 0.01 to about12 hours, or from about 0.01 to about 6 hours, or from about 0.01 toabout 3 hours, or from about 0.01 to about 1 hour, or from about 0.01 toabout 0.5 hour, or from about 0.1 to about 48 hours, or from about 0.1to about 36 hours, or from about 0.1 to about 24 hours, or from about0.1 to about 12 hours, or from about 0.1 to about 6 hours, or from about0.1 to about 3 hours, or from about 0.1 to about 3 hours, or from about0.1 to about 1 hour, or from about 0.1 to about 0.5 hour at atemperature in the range of about 150 to about 300° C., or about 200° C.to about 250° C., and then restarting the flow of synthesis gas into thereactor. The purge may be at a flow rate lower than or equal to orgreater than the synthesis gas feed rate to the reactor prior tostoppage.

In an embodiment, the flow of synthesis gas into the reactor 110 isstopped, and the Fischer-Tropsch process may be restarted by a methodcomprising: maintaining the reactor at the pre-stop operatingtemperature for a period of time of up to about 48 hours, or up to about36 hours, or up to about 24 hours, or up to about 36 hours, or up toabout 12 hours, or up to about 6 hours, or up to about 3 hours, or up toabout 1 hour, or from about 0.01 to about 48 hours, or from about 0.01to about 36 hours, or from about 0.01 to about 24 hours, or from about0.01 to about 12 hours, or from about 0.01 to about 6 hours, or fromabout 0.01 to about 2 hours, or from about 0.01 to about 1 hour, or fromabout 0.01 to about 0.5 hour, or from about 0.1 to about 48 hours, orfrom about 0.1 to about 36 hours, or from about 0.1 to about 24 hours,or from about 0.1 to about 12 hours, or from about 0.1 to about 6 hours,or from about 0.1 to about 3 hours, or from about 0.1 to about 2 hours,or from about 0.1 to about 1 hour, or from about 0.1 to about 0.5 hour,then flowing hydrogen, nitrogen or desulfurized natural gas into thereactor to purge the reactor of reactants and Fischer-Tropsch product,holding the reactor in a purge gas environment for a period of time ofup to about 48 hours, or up to about 36 hours, or up to about 24 hours,or up to about 12 hours, or up to about 6 hours, or up to about 3 hours,or up to about 1 hour, cooling the reactor to a temperature lower thanthe operating temperature (e.g., in the range of about 25 to about 150°C.); and then restarting the flow of synthesis gas into the reactor. Thepurge may be at a flow rate lower than or equal to or greater than thesynthesis gas feed rate to the reactor prior to stoppage. The reactortemperature may then be ramped up to the operating temperature that wasused prior to the stop within a short period of time by heating at arate of up to about 5° C. per hour, or up to about 10° C. per hour, orup to about 15° C. per hour, or up to about 30° C. per hour, or up toabout 60° C. per hour.

In an embodiment, the flow of synthesis gas into the reactor 110 isstopped, and the Fischer-Tropsch process may be restarted by a methodcomprising: maintaining the reactor at the pre-stop operatingtemperature and pressure for a period of up to about 48 hours, or up toabout 36 hours, or up to about 24 hours, or up to about 12 hours, or upto about 6 hours, or up to about 3 hours, or up to about 1 hour, or fromabout 0.01 to about 48 hours, or from about 0.01 to about 36 hours, orfrom about 0.01 to about 24 hours, or from about 0.01 to about 12 hours,or from about 0.01 to about 6 hours, or from about 0.01 to about 3hours, or from about 0.01 to about 1 hour, or from about 0.01 to about0.5 hour, or from about 0.1 to about 48 hours, or from about 0.1 toabout 36 hours, or from about 0.1 to about 24 hours, or from about 0.1to about 12 hours, or from about 0.1 to about 6 hours, or from about 0.1to about 3 hours, or from about 0.1 to about 2 hours, or from about 0.1to about 1 hour, or from about 0.1 to about 0.5 hour then depressurizingthe reactor to a pressure (lower than the operating pressure) of up toabout 20 atmospheres, or up to about 10 atmospheres, or up to about 5atmospheres, and then flowing hydrogen, nitrogen, or desulfurizednatural gas into the reactor to purge the reactor of reactants andeffluent, holding the catalyst reactor in a purge gas environment for aperiod of time of up to about 48 hours, or up to about 36 hours, up toabout 24 hours, or up to about 12 hours, or up to about 6 hours, or upto about 3 hours, or up to about 1 hour at a temperature in the rangefrom about 25° C. to about 300° C., or about 150° C. to about 250° C.;and then restarting the flow of synthesis gas into the reactor andpressurizing the reactor to the target operating pressure. The purge maybe at a flow rate lower than or equal to or greater than the synthesisgas feed rate to the reactor prior to stoppage. If the reactortemperature is reduced to below the pre-stop operating temperature, thereactor temperature may then be ramped up to the operating temperatureprior to the stop within a short time by heating at a rate as high of upto about 5° C. per hour, or up to about 10° C. per hour, or up to about15° C. per hour, or up to about 30° C. per hour, or up to about 60° C.per hour.

In an embodiment, the flow of synthesis gas into the reactor 110 isstopped, and the Fischer-Tropsch process is restarted by a methodcomprising: restarting the flow of synthesis gas into the reactor, thetemperature of the synthesis gas flowing into the reactor being lessthan up to about 10° C. of the desired reaction temperature in thereactor, or less than up to about 5° C. of the desired reactiontemperature. The reaction temperature may be controlled with a coolantflowing in a heat exchanger in thermal contact with the reactor, andduring the step of restarting the flow of synthesis gas into thereactor, the temperature of the coolant in the heat exchanger may beless than up to about 10° C., or less than up to about 5° C., of thereaction temperature in the reactor. In this embodiment, theFischer-Tropsch catalyst may comprise a wet catalyst.

In an embodiment, the flow of synthesis gas into the reactor 110 isstopped, and the Fischer-Tropsch process is restarted by a methodcomprising: rejuvenating or regenerating the catalyst; and restartingthe flow of synthesis gas into the reactor, the temperature of thesynthesis gas flowing into the reactor being at the desired reactiontemperature. The reaction temperature in the reactor may be controlledwith a heat exchange fluid flowing in a heat exchanger in thermalcontact with the reactor, and during the restarting of the flow ofsynthesis gas into the reactor and the flow of effluent out of thereactor, the temperature of the heat exchange fluid in the heatexchanger may be lower than the reaction temperature in the reactor, forexample, up to about 10° C. lower, or up to about 5° C. lower, than thereaction temperature in the reactor.

In an embodiment, the flow of synthesis gas into and the flow ofeffluent out of the reactor 110 is stopped, and the catalyst is a wetcatalyst containing a wax. The Fischer-Tropsch process may be restartedby a method comprising: rejuvenating the catalyst by flowing hydrogen incontact with the catalyst at a temperature of up to about 400° C. andthen restarting the flow of synthesis gas into the reactor. The reactiontemperature in the reactor may be controlled with a heat exchange fluidin thermal contact with the reactor, and during the restarting of theflow of synthesis gas into the reactor and the flow of effluent out ofthe reactor, the temperature of the heat exchange fluid may be lowerthan the reaction temperature in the reactor, for example, up to about10° C. lower, or up to about 5° C. lower, than the reaction temperaturein the reactor.

In an embodiment, the flow of synthesis gas into and the flow ofeffluent out of the reactor 110 is stopped, and the catalyst is a wetcatalyst containing a wax. The Fischer-Tropsch process may be restartedby a method comprising: regenerating the catalysts by removing wax fromthe catalyst by flowing hydrogen in contact with the catalyst at atemperature of up to about 350° C., or up to about 400° C., or up toabout 450° C. This may be followed by oxidizing the catalyst in air(21%02) flowing air into the reactor in contact with the catalyst at atemperature in the range of up to about 200° C., or up to about 300° C.,or up to about 350° C., for a period of time in the range from about 1to about 12 hours, or even with no hold time at the maximum temperature.This may be followed by another reduction step where hydrogen flows incontact with the catalyst at a temperature up to about 400° C., or up toabout 350° C. The reactor is then cooled to the desired reactiontemperature. The flow of synthesis gas into the reactor in contact withthe catalyst may then be commenced. The reaction temperature in thereactor may be controlled with a heat exchange fluid in thermal contactwith the reactor, and during the restarting of the flow of synthesis gasinto the reactor, the temperature of the heat exchange fluid may belower than the reaction temperature in the reactor, for example, up toabout 10° C. lower, or up to about 5° C. lower, than the reactiontemperature in the reactor.

Referring to FIG. 2 , the synthesis gas conversion (e.g.,Fischer-Tropsch process) may be conducted in a plant 1000 comprising aplurality of reaction trains 100 a, 100 b and 100 c. Each reaction train100 a, 100 b and 100 c comprises a Fischer-Tropsch reactor 110 a, 110 b,110 c, respectively, containing a Fischer-Tropsch catalyst. The reactiontrains 100 a, 100 b and 100 c may be connected to a common reactant feedstream 1010 comprising fresh synthesis gas. The flow of synthesis gasinto the Fischer-Tropsch reactor of one or more of the reaction trains(100 a, 100 b, 100 c) in the plant 1000 may be stopped, while the flowof synthesis gas into and the flow of effluent out of the remainder ofthe reaction trains (100 a, 100 b, 100 c) in the plant 1000 iscontinued. The method may comprise: (A) flowing the reactant feed stream1010 at an overall desired process flow rate to the plurality ofreaction trains (100 a, 100 b, 100 c) in the plant 1000; (B) dividingthe reactant feed stream into a plurality of reactant substreams (120 a,120 b, 120 c); (C) flowing each reactant substream through a separatereaction train to convert the reactants in the reactant substream to aFischer-Tropsch product; (D) stopping the flow of a reactant substream(120 a, 120 b, 120 c) to one or more of the reaction trains (100 a, 100b, 100 c); and (E) continuing to flow the reactant feed 1010 stream tothe remainder of reaction trains (100 a, 100 b, 100 c) in the plant 1000at an overall process flow rate for the flow of fresh synthesis gas thatis within up to about 15%, or up to about 10%, or up to about 5%, or upto about 2%, or up to about 1%, of the overall process flow rate for theflow of fresh synthesis gas used in step (A). During step (C) a mixtureof fresh synthesis gas and a recycled tail gas may flow into theFischer-Tropsch reactor of each reaction train. During step (E) the flowof recycled tail gas into the Fischer-Tropsch reactor of the one or moreof the remainder of the reaction trains in the plant may be stopped orreduced in order to allow for maintaining a desired overall flow offresh synthesis gas into the plant.

Referring to FIG. 2 , in an embodiment, the synthesis gas conversionprocess (e.g., Fischer-Tropsch process) is conducted in plant 1000comprising a plurality of reaction trains 100 a, 100 b and 100 c. Eachreaction train comprises a Fischer-Tropsch reactor 110 a, 110 b, and 110c containing a Fischer-Tropsch catalyst. The reaction trains 100 a, 100b, and 100 c are connected to a reactant feed stream 1010 comprisingfresh synthesis gas. The flow of synthesis gas into and the flow ofeffluent out of the Fischer-Tropsch reactor of one or more of thereaction trains (100 a, 100 b, 100 c) may be stopped, while the flow ofsynthesis gas into and the flow of effluent out of the other reactiontrains in the plant 1000 may be continued. The method may comprise:flowing the reactant feed stream to the plurality of reaction trains(100 a, 100 b, 100 c); dividing the reactant feed stream 1010 into areactant substream (120 a, 120 b, 120 c) for each reaction train (100 a,100 b, 100 c); and flowing each reactant substream through a reactiontrain to convert the reactants in the reactant substream to aFischer-Tropsch product. The temperature of the Fischer-Tropsch productflowing out of each reaction train in the plant 1000 may be the same orsubstantially the same. In an embodiment, the temperature of theFischer-Tropsch product flowing out of one of the reaction trains may bewithin about 10° C., or about 5° C., or about 2° C., or about 1° C., ofthe temperature of the Fischer-Tropsch product flowing out of another ofthe reaction trains in the plant 1000.

When a microchannel reactor is used, the fresh synthesis gas or thereactant mixture flows through one or more process microchannels in thereactor in contact with a Fischer-Tropsch catalyst to form aFischer-Tropsch product. The Fischer-Tropsch product may comprise amixture of FT liquid, FT wax, and FT tail gas. The FT tail gas maycomprise CO and H₂. The reaction is exothermic. The reaction may becontrolled using a heat exchange fluid which flows through theFischer-Tropsch reactor 110 as indicated by arrows 170 and 180. In anembodiment, the heat exchange fluid may comprise steam. TheFischer-Tropsch product flows out of the Fischer-Tropsch reactor 110 asindicated by arrow 150. FT tail gas is separated from theFischer-Tropsch product, as indicated by arrow 130, and recycled to becombined with the fresh synthesis gas. Part of the FT tail gas may beseparated from the process, as indicated by arrow 135, if it is desiredto adjust the ratio of fresh synthesis gas to FT tail gas in thereactant mixture. With FT tail gas separated from the Fischer-Tropschproduct, the remainder of the Fischer-Tropsch product, which isindicated by arrow 160, comprises a FT liquid and/or FT wax, which canbe subjected to further processing.

The synthesis gas conversion (e.g., Fischer-Tropsch reaction process)illustrated in FIG. 1 may be referred to as a reaction train. In aproduction facility or plant, the Fischer-Tropsch reaction may beconducted in a plurality of reaction trains that may be connected to acommon source of fresh synthesis gas. Referring to FIG. 2 , plant 1000contains three reaction trains 100 a, 100 b and 100 c, each of which isconnected to fresh synthesis gas line 1010. Although three reactiontrains are illustrated in FIG. 2 , it is to be understood that plant1000 may contain any desired number of reaction trains, for example,from 1 to about 1000 reaction trains, or from 1 to about 500 reactiontrains, or from 1 to about 200 reaction trains, or from 1 to about 100reaction trains, or from 1 to about 50 reaction trains, or from 1 toabout 20 reaction trains, or from 1 to about 10 reaction trains, or from1 to about 5 reaction trains. In operation, fresh synthesis gas fromfresh synthesis gas line 1010 is divided into substreams 120 a, 120 band 120 c, which flow into reaction trains 100 a, 100 b and 100 c,respectively.

The reaction train 100 a employs the use of Fischer-Tropsch reactor 110a. In operation, fresh synthesis gas substream 120 a flows from freshsynthesis gas line 1010 into the Fischer-Tropsch reactor 110 a.Alternatively, the fresh synthesis gas may be combined with recycledtail gas 130 a to form reactant mixture 140 a which flows into theFischer-Tropsch reactor 110 a. The fresh synthesis gas may be combinedwith the recycled tail gas upstream of the Fischer-Tropsch reactor 110a, as shown in FIG. 2 , or in the Fischer-Tropsch reactor 110 a. Thereaction may be controlled using a heat exchange fluid which flowsthrough the Fischer-Tropsch reactor 110 a as indicated by arrows 170 aand 180 a. The Fischer-Tropsch product flows out of the Fischer-Tropschreactor 110 a as indicated by arrow 150 a. FT tail gas is separated fromthe Fischer-Tropsch product, as indicated by arrow 130 a, and may berecycled to be combined with the fresh synthesis gas. Part of the FTtail gas may be separated from the process, as indicated by arrow 135 a,if it is desired to adjust the ratio of fresh synthesis gas to FT tailgas in the reactant mixture. With FT tail gas separated from theFischer-Tropsch product, the remainder of the Fischer-Tropsch product,which is indicated by arrow 160 a, comprises a FT liquid and/or FT wax.The FT liquid and/or FT wax can be subjected to further processing.

The reaction train 100 b employs the use of Fischer-Tropsch reactor 110b. In operation, fresh synthesis gas substream 120 b flows from freshsynthesis gas line 1010 into the Fischer-Tropsch reactor 110 b.Alternatively, the fresh synthesis gas may be combined with recycledtail gas 130 b to form reactant mixture 140 b which flows into theFischer-Tropsch reactor 110 b. The fresh synthesis gas may be combinedwith the recycled tail gas upstream of the Fischer-Tropsch reactor 110b, as shown in FIG. 2 , or in the Fischer-Tropsch reactor 110 b. Thereaction may be controlled using a heat exchange fluid which flowsthrough the Fischer-Tropsch reactor 110 b as indicated by arrows 170 band 180 b. The Fischer-Tropsch product flows out of the Fischer-Tropschreactor 110 b as indicated by arrow 150 b. FT tail gas is separated fromthe Fischer-Tropsch product, as indicated by arrow 130 b, and may berecycled to be combined with the fresh synthesis gas. Part of the FTtail gas may be separated from the process, as indicated by arrow 135 b,if it is desired to adjust the ratio of fresh synthesis gas to FT tailgas in the reactant mixture. With FT tail gas separated from theFischer-Tropsch product, the remainder of the Fischer-Tropsch product,which is indicated by arrow 160 b, comprises a FT liquid and/or FT wax.The FT liquid and/or FT wax can be subjected to further processing.

The reaction train 100 c employs the use of Fischer-Tropsch reactor 110c. In operation, fresh synthesis gas substream 120 c flows from freshsynthesis gas line 1010 into the Fischer-Tropsch reactor 110 c.Alternatively, the fresh synthesis gas may be combined with recycledtail gas 130 c to form reactant mixture 140 c which flows into theFischer-Tropsch reactor 110 c. The fresh synthesis gas may be combinedwith the recycled tail gas upstream of the Fischer-Tropsch reactor 110c, as shown in FIG. 2 , or in the Fischer-Tropsch reactor 110 c. Thereaction may be controlled using a heat exchange fluid which flowsthrough the Fischer-Tropsch reactor 110 c as indicated by arrows 170 cand 180 c. The Fischer-Tropsch product flows out of the Fischer-Tropschreactor 110 c as indicated by arrow 150 c. FT tail gas is separated fromthe Fischer-Tropsch product, as indicated by arrow 130 c, and may berecycled to be combined with the fresh synthesis gas. Part of the FTtail gas may be separated from the process, as indicated by arrow 135 c,if it is desired to adjust the ratio of fresh synthesis gas to FT tailgas in the reactant mixture. With FT tail gas separated from theFischer-Tropsch product, the remainder of the Fischer-Tropsch product,which is indicated by arrow 160 c, comprises a FT liquid and/or FT wax.The FT liquid and/or FT wax can be subjected to further processing.

In an embodiment, the temperature of the product streams 150 a, 150 band 150 c may be the same, or substantially the same, that is, withinabout 10° C., or about 5° C., or about 2° C., or about 1° C. of eachother.

The synthesis gas conversion microchannel reactor (e.g., Fischer-Tropschmicrochannel reactor) that may be used is illustrated in FIG. 3 .Referring to FIG. 3 , microchannel reactor 200 comprises containmentvessel 210 which contains or houses three microchannel reactor cores220. Although FIG. 3 shows containment vessel 210 containing or housingthree microchannel reactor cores 220, it is to be understood that anydesired number of microchannel reactor cores 220 may be contained orhoused within containment vessel 210. For example, containment vessel210 may be used to contain or house from 1 to about 12 microchannelreactor cores, or from 1 to about 8 microchannel reactor cores, or from1 to about 4 microchannel reactor cores. The containment vessel 210 maybe a pressurizable vessel. The containment vessel 210 includes inletsand outlets 230 allowing for the flow of reactants into the microchannelreactor cores 220, product out of the microchannel reactor cores 220,and heat exchange fluid into and out of the microchannel reactor cores220.

One of the inlets 230 may be connected to a header or manifold which isprovided for flowing reactants to process microchannels in each of themicrochannel reactor cores 220. One of the inlets 230 is connected to aheader or manifold which is provided for flowing a heat exchange fluidto heat exchange channels in each of the microchannel reactor cores 220.One of the outlets 230 is connected to a manifold or footer whichprovides for product flowing out of the process microchannels in each ofthe microchannel reactor cores 220. One of the outlets 230 is connectedto a manifold or footer to provide for the flow of the heat exchangefluid out of the heat exchange channels in each of the microchannelreactor cores 220.

The containment vessel 200 may be constructed using any suitablematerial sufficient for countering operating pressures that may developwithin the microchannel reactor cores 220. For example, the shell 240and reinforcing ribs 242 of the containment vessel 210 may beconstructed of cast steel. The flanges 245, couplings and pipes may beconstructed of 316 stainless steel. The containment vessel 210 may haveany desired diameter, for example, from about 10 to about 1000 cm, orfrom about 50 to about 300 cm. The axial length of the containmentvessel 210 may be of any desired value, for example, from about 0.5 toabout 50 meters, or from about 1 to about 20 meters, or from about 1 toabout 10 meters, or from about 1 to about 5 meters.

The microchannel reactor core 220 may comprise one or more layers ofprocess microchannels and heat exchange channels stacked one above theother or positioned side-by-side to form a microchannel reactor core;see, FIGS. 4 and 5 . The microchannel reactor core 220 may have the formof a three-dimensional block which has six faces that are squares orrectangles. The microchannel reactor core 220 may have the samecross-section along a length. The microchannel reactor core 220 may bein the form of a parallel or cubic block or prism. The microchannelreactor core 220 may have a length, width and height of any dimension,for example, a length in the range from about 0.1 to about 5 meters,about 0.01 to about 3 meters, or from about 0.1 to about 2 meters, orfrom about 0.1 to about 1.5 meters, or from about 0.1 to about 1 meter,or from about 0.1 to about 0.75 meter, or from about 0.1 to about 0.5meter, or from about 0.1 to about 0.25 meter, or from about 0.1 to about0.15 meter. The microchannel core 220 may have a width in the range fromabout 0.1 to about 5 meters, about 0.1 to about 3 meters, or from about0.1 to about 2 meters, or from about 0.1 to about 1.5 meters, or fromabout 0.1 to about 1 meter, or from about 0.1 to about 0.75 meter, orfrom about 0.1 to about 0.5 meter, or from about 0.1 to about 0.25meter, or from about 0.1 to about 0.15 meter. The microchannel reactorcore 220 may have a height in the range from about 0.1 to about 5meters, or about 0.1 to about 3 meters, or from about 0.1 to about 2meters, or from about 0.1 to about 1.5 meters, or from about 0.1 toabout 1 meter, or from about 0.1 to about 0.75 meter, or from about 0.1to about 0.5 meter, or from about 0.1 to about 0.25 meter, or from about0.1 to about 0.15 meter.

The microchannel reactor core 220 may contain a plurality of layers ofprocess microchannels and a plurality of layers of heat exchangechannels. These are illustrated in FIGS. 5-7 . Referring to FIGS. 5-7 ,microchannel reactor core 220 may contain from 1 to about 1000 layers300 of process microchannels 310, or from about 1 to about 500, or from1 to about 400, or from 1 to about 300, or from 1 to about 200, or from1 to about 100, or from 1 to about 50, or from 1 to about 25, or from 1to about 10, or from 1 to about 5 of such layers. The catalyst may bepositioned in the process microchannels 310 and may be in any formincluding fixed beds of particulate solids or any of the variousstructured catalyst forms described below. The microchannel reactor core220 may contain from 1 to about 1000 layers 350 of heat exchangechannels 355, or from 1 to about 500, or from 1 to about 400, or from 1to about 300, or from 1 to about 200, or from 1 to about 100, or from 1to about 50, or from 1 to about 25, or from 1 to about 10, or from 1 toabout 5 of such layers.

The layers 300 of process microchannels 310 may be constructed usingwaveforms in the form of right angled corrugated sheet inserts 315.These waveform inserts 315 may have rounded edges rather than sharpedges. These waveform inserts may have a sinusoidel form as shown inFIGS. 5 and 6 . The waveform insert 315 may be positioned betweenopposing planar plates 316 and 317. This is shown in FIGS. 5 and 6 . Inthis manner the microchannels may be defined on three sides by thewaveform insert 315 and on the fourth side by one of the planar plates316 or 317. The layers of process microchannels as well as the layers ofheat exchange channels may be formed in this manner. Also, the spacebetween opposed planar plates 316 and 317 containing the waveform 315may have the dimensions of a microchannel (e.g., a height of up to about10 mm) and may therefore also be referred to as a microchannel.Microchannel reactors made using waveforms are disclosed in WO2008/030467, which is incorporated herein by reference.

Each layer 300 of process microchannels 310 may have a height (h) in therange from about 0.1 to about 10 mm. Each layer 300 may have a width (w)in the range from about 0.1 to about 5 meters, or about 0.1 to about 3meters, or about 0.1 to about 2 meters, or about 0.1 to about 1 meter,or about 0.1 to about 0.5 meter. The length (l) of each layer 300 may beup to about 5 meters, or from about 0.1 to about 5 meters, or from about0.1 to about 3 meters, or from about 0.1 to about 2.5 meters, or fromabout 0.1 to about 2 meters, or from about 0.1 to about 1.5 meters, orfrom about 0.1 to about 1 meter, or from about 0.15 to about 5 meters,or from about 0.15 to about 5 meters, or from about 0.15 to about 3meters, or from about 0.15 to about 2.5 meters, or from about 0.15 toabout 2 meters, or from about 0.15 to about 1.5 meters, or from about0.15 to about 1 meter. The length (l) may be in the range from about 0.1to about 0.8 meter, or from about 0.1 to about 0.6 meter, or from about0.1 to about 0.5 meter, or from about 0.1 to about 0.3 meter.

Each layer 350 of heat exchange channels 355 may have a height (h) inthe range from about 0.1 to about 10 mm. Each layer 350 may have a width(w) in the range from 0.1 to about 5 meters, or about 0.1 to about 3meters, or about 0.1 to about 2 meters, or about 0.1 to about 1 meter,or about 0.01 to about 0.5 meter. The length (l) of each layer 350 maybe from about 0.1 to about 5 meters, or from about 0.1 to about 3meters, or from about 0.1 to about 2.5 meters, or from about 0.1 toabout 2 meters, or from about 0.1 to about 1.5 meters, or from about 0.1to about 1 meter, or from about 0.15 to about 3 meters, or from about0.15 to about 2.5 meters, or from about 0.15 to about 2 meters, or fromabout 0.15 to about 1.5 meters, or from about 0.15 to about 1 meter. Thelength (l) may be in the range from about 0.1 to about 0.8 meter, orfrom about 0.1 to about 0.6 meter, or from about 0.05 to about 0.5meter, or from about 0.1 to about 0.3 meter.

Each layer 300 of process microchannels 310 may have from 1 to about5000 process microchannels 310, or from 1 to about 2000 processmicrochannels, or from 1 to about 1000 process microchannels, or from 1to about 500 process microchannels, or from 1 to about 250 processmicrochannels, or from 1 to about 100 process microchannels, or from 1to about 50 process microchannels. The process microchannels 310 mayhave cross sections having any shape, for example, square, rectangle,circle, semi-circle, etc. The internal height of each processmicrochannel 310 may be considered to be the smaller of the internaldimensions normal to the direction of flow of reactants and productthrough the process microchannel. Each of the process microchannels 310may have internal height or width in the range of up to about 10 mm, orfrom about 0.05 to about 10 mm, or from about 0.05 to about 8 mm, orfrom about 0.05 to about 7 mm, or from about 0.05 to about 5 mm, or fromabout 0.05 to about 3 mm, or from about 0.05 to about 2 mm, or fromabout 0.05 to about 1.5 mm, or from about 1 to about 10 mm, or fromabout 1 to about 8 mm, or from about 1 to about 7 mm, or from about 1 toabout 5 mm, or from about 1 to about 3 mm, or from about 1 to about 2mm, or from about 1 to about 1.5 mm. The other internal dimension ofheight or width may be of any dimension, for example, up to about 5meters, or about 0.001 to about 5 meters, or about 0.001 to about 3meters, or about 0.001 to about 2 meters, or about 0.001 to about 1meter, or about 0.01 to about 0.5 meter, or about 1 to about 10 mm, orabout 1 to about 8 mm, or about 1 to about 7 mm, or about 1 to about 5mm, or about 1 to about 3 mm, or about 1 to about 2 mm. The length maybe of any dimension, for example, up to about 5 meters, or from about0.1 to about 5 meters, or from about 0.1 to about 3 meters, or fromabout 0.1 to about 2.5 meters, or from about 0.1 to about 2 meters, orfrom about 0.1 to about 1.5 meters, or from about 0.1 to about 1 meter,or from about 0.15 to about 5 meters, or from about 0.15 to about 3meters, or from about 0.15 to about 2.5 meters, or from about 0.15 toabout 2 meters, or from about 0.15 to about 1.5 meters, or from about0.15 to about 1 meter. The length may be in the range from about 0.1 toabout 0.8 meter, or from about 0.1 to about 0.6 meter, or from about 0.1to about 0.5 meter, or from about 0.1 to about 0.3 meter.

Each layer 350 of heat exchange channels 355 may have from 1 to about5000 heat exchange channels, or from 1 to about 2000 heat exchangechannels, or from 1 to about 1000 heat exchange channels, or from 1 toabout 500 heat exchange channels, or from 1 to about 250 heat exchangechannels, or from 1 to about 100 heat exchange channels, or from 1 toabout 50 heat exchange channels. The heat exchange channels 355 may bemicrochannels or they may have larger dimensions that would classifythem as not being microchannels. Each of the heat exchange channels 355may have a cross section having any shape, for example, a square,rectangle, circle, semi-circle, etc. The internal height of each heatexchange channel 355 may be considered to be the smaller of the internaldimensions normal to the direction of flow of heat exchange fluid in theheat exchange channels. Each of the heat exchange channels 355 may haveinternal height or width in the range of up to about 10 mm, or fromabout 0.05 to about 10 mm, or from about 0.05 to about 8 mm, or fromabout 0.05 to about 7 mm, or from about 0.05 to about 5 mm, or fromabout 0.05 to about 3 mm, or from about 0.05 to about 2 mm, or fromabout 0.05 to about 1.5 mm, or from about 1 to about 10 mm, or fromabout 1 to about 8 mm, or from about 1 to about 7 mm, or from about 1 toabout 5 mm, or from about 1 to about 3 mm, or from about 1 to about 2mm, or from about 1 to about 1.5 mm. The other internal dimension ofheight or width may be of any dimension, for example, up to about 5meters, or about 0.001 to about 5 meters, or about 0.001 to about 3meters, or about 0.001 to about 2 meters, or about 0.001 to about 1meter, or about 0.01 to about 0.5 meter, or about 1 to about 10 mm, orabout 1 to about 8 mm, or about 1 to about 7 mm, or about 1 to about 5mm, or about 1 to about 3 mm, or about 1 to about 2 mm. The length maybe of any dimension, for example, up to about 5 meters, or about 0.1 toabout 5 meters, or from about 0.1 to about 3 meters, or from about 0.1to about 2.5 meters, or from about 0.1 to about 2 meters, or from about0.1 to about 1.5 meters, or from about 0.1 to about 1 meter, or fromabout 0.15 to about 5 meters, or from about 0.15 to about 3 meters, orfrom about 0.15 to about 2.5 meters, or from about 0.15 to about 2meters, or from about 0.15 to about 1.5 meters, or from about 0.15 toabout 1 meter. The length may be in the range from about 0.1 to about0.8 meter, or from about 0.1 to about 0.6 meter, or from about 0.05 toabout 0.5 meter, or from about 0.1 to about 0.3 meter.

The microchannel reactor core 220 may be made of any material thatprovides sufficient strength, dimensional stability and heat transfercharacteristics to permit operation of the desired process. Thesematerials may include aluminum; titanium; nickel; platinum; rhodium;copper; chromium; alloys of any of the foregoing metals; brass; steel(e.g., stainless steel); quartz; silicon; or a combination of two ormore thereof. Each microchannel reactor may be constructed of stainlesssteel with one or more copper or aluminum waveforms being used forforming the channels.

The microchannel reactor core 220 may be fabricated using knowntechniques including wire electrodischarge machining, conventionalmachining, laser cutting, photochemical machining, electrochemicalmachining, molding, water jet, stamping, etching (for example, chemical,photochemical or plasma etching) and combinations thereof.

The microchannel reactor core 220 may be constructed by forming plateswith portions removed that allow flow passage. A stack of plates may beassembled via diffusion bonding, laser welding, diffusion brazing, andsimilar methods to form an integrated device. The microchannel reactorsmay be assembled using a combination of plates and partial plates orstrips. In this method, the channels or void areas may be formed byassembling strips or partial plates to reduce the amount of materialrequired.

The microchannel reactor core 220 may comprise a plurality of plates ina stack defining a plurality of process layers and a plurality of heatexchange layers, each plate having a peripheral edge, the peripheraledge of each plate or shim being welded to the peripheral edge of thenext adjacent plate to provide a perimeter seal for the stack. This isshown in US 2012/0095268 A1, which is incorporated herein by reference.

The containment vessel 210 may include a control mechanism to maintainthe pressure within the containment vessel at a level that is at leastas high as the internal pressure within the microchannel reactor cores220. The internal pressure within the containment vessel 210 may be inthe range from about 10 to about 60 atmospheres, or from about 15 toabout 30 atmospheres during the operation of a synthesis gas conversionprocess (e.g., Fischer-Tropsch process). The control mechanism formaintaining pressure within the containment vessel may comprise a checkvalve and/or a pressure regulator. The check valve or regulator may beprogrammed to activate at any desired internal pressure for thecontainment vessel. Either or both of these may be used in combinationwith a system of pipes, valves, controllers, and the like, to ensurethat the pressure in the containment vessel 210 is maintained at a levelthat is at least as high as the internal pressure within themicrochannel reactor cores 220. This is done in part to protect weldsused to form the microchannel cores 220. A significant decrease in thepressure within the containment vessel 210 without a correspondingdecrease of the internal pressure within the microchannel reactor cores220 could result in a costly rupture of the welds within themicrochannel reactor cores 220. The control mechanism may be designed toallow for diversion of one or more process gases into the containmentvessel in the event the pressure exerted by the containment gasdecreases.

In the design of a microchannel reactor it may be advantageous toprovide a tailored heat exchange profile along the length of the processmicrochannels in order to optimize the reaction. This may beaccomplished by matching the local release of heat given off by thesynthesis gas conversion reaction (e.g., Fischer-Tropsch reaction)conducted in the process microchannels with heat removal or coolingprovided by heat exchange fluid in heat exchange channels in themicrochannel reactor. The extent of the synthesis gas conversionreaction (e.g., Fischer-Tropsch reaction) and the consequent heatrelease provided by the reaction may be higher in the front or upstreamsections of the reaction zones in the process microchannels as comparedto the back or downstream sections of the reaction zones. Consequently,the matching cooling requirements may be higher in the upstream sectionof the reaction zones as compared to the downstream sections of thereaction zones. Tailored heat exchange may be accomplished by providingmore heat exchange or cooling channels, and consequently the flow ofmore heat exchange or cooling fluid, in thermal contact with upstreamsections of the reaction zones in the process microchannels as comparedto the downstream sections of the reaction zones. Alternatively oradditionally, a tailored heat exchange profile may be provided byvarying the flow rate of heat exchange fluid in the heat exchangechannels. In areas where additional heat exchange or cooling is desired,the flow rate of the heat exchange fluid may be increased as compared toareas where less heat exchange or cooling is required. For example, ahigher rate of flow of heat exchange fluid may be advantageous in theheat exchange channels in thermal contact with the upstream sections ofthe reaction zones in the process microchannels as compared to the heatexchange channels in thermal contact with the downstream sections of thereaction zones. Heat transfer from the process microchannels to the heatexchange channels may be designed for optimum performance by selectingoptimum heat exchange channel dimensions and/or the rate of flow of heatexchange fluid per individual or groups of heat exchange channels.Additional design alternatives for tailoring heat exchange may relate tothe selection and design of the synthesis gas conversion catalyst (e.g.,Fischer-Tropsch catalyst) such as, particle size, catalyst formulation,packing density, use of a graded catalyst, or other chemical or physicalcharacteristics, at specific locations within the process microchannels.These design alternatives may impact both heat release from the processmicrochannels as well as heat transfer to the heat exchange fluid.Temperature differentials between the process microchannels and the heatexchange channels, which may provide a driving force for heat transfer,may be constant or may vary along the length of the processmicrochannels.

The process microchannels may contain one or more surface features inthe form of depressions in and/or projections from one or more interiorwalls of the process microchannels. The surface features may be used todisrupt the flow of fluid flowing in the channels. These disruptions inflow may enhance mixing and/or heat transfer. The surface features maybe in the form of patterned surfaces. The microchannel reactor core 220may be made by laminating a plurality of plates together. One or bothmajor surfaces of the plates may contain surface features.Alternatively, the microchannel reactor core 220 may be assembled usingsome plates and some strips, or partial plates to reduce the totalamount of metal required to construct the device. A plate containingsurface features may be paired (on opposite sides of a microchannel)with another plate containing surface features. Pairing may createbetter mixing or heat transfer enhancement as compared to channels withsurface features on only one major surface. The patterning may comprisediagonal recesses that are disposed over substantially the entire widthof a microchannel surface. The patterned surface feature area of a wallmay occupy part of or the entire length of a microchannel surface.Surface features may be positioned over at least about 10%, or at leastabout 20%, or at least about 50%, or at least about 80% of the length ofa channel surface. Each diagonal recesses may comprise one or moreangles relative to the flow direction. Successive recessed surfacefeatures may comprise similar or alternate angles relative to otherrecessed surface features.

The synthesis gas conversion process (e.g., Fischer-Tropsch processmicrochannels) may be characterized by having bulk flow paths. The term“bulk flow path” refers to an open path (contiguous bulk flow region)within the process microchannels or combustion channel. A contiguousbulk flow region allows rapid fluid flow through the channels withoutlarge pressure drops. In one embodiment, the flow of fluid in the bulkflow region is laminar. Bulk flow regions within each processmicrochannel or combustion channel may have a cross-sectional area ofabout 0.05 to about 10,000 mm², or about 0.05 to about 5000 mm², orabout 0.1 to about 2500 mm². The bulk flow regions may comprise fromabout 5% to about 95%, or about 30% to about 80% of the cross-section ofthe process microchannels or combustion channel.

The contact time of the reactants with the catalyst may range up toabout 3600 milliseconds (ms), or up to about 2000 ms, or in the rangefrom about 10 to about 2600 ms, or from about 10 ms to about 2000 ms, orabout 20 ms to about 500 ms, or from about 200 to about 400 ms, or fromabout 240 to about 350 ms.

The space velocity (or gas hourly space velocity (GHSV)) for the flow offluid in the process microchannels may be at least about 1000 hr⁻¹(normal liters of feed/hour/liter of volume within the processmicrochannels), or at least about 1800 hr⁻¹, or from about 1000 to about1,000,000 hr⁻¹, or from about 5000 to about 20,000 hr⁻¹.

The pressure within the process microchannels may be up to about 100atmospheres, or in the range from about 1 to about 100 atmospheres, orfrom about 1 to about 75 atmospheres, or from about 2 to about 40atmospheres, or from about 2 to about 10 atmospheres, or from about 10to about 50 atmospheres, or from about 20 to about 30 atmospheres.

The pressure drop of fluids as they flow in the process microchannelsmay range up to about 30 atmospheres per meter of length of channel(atm/m), or up to about 25 atm/m, or up to about 20 atm/m. The pressuredrop may be in the range from about 10 to about 20 atm/m.

The Reynolds Number for the flow of fluid in the process microchannelsmay be in the range of about 10 to about 4000, or about 100 to about2000.

The average temperature in the process microchannels may be in the rangefrom about 150 to about 300° C., or in the range from about 175 to about225° C., of in the range from about 190 to about 220° C., or from about195 to about 215° C.

The heat exchange fluid entering the heat exchange channels of themicrochannel reactor core 220 may be at a temperature in the range ofabout 100° C. to about 400° C., or about 200° C. to about 300° C. Theheat exchange fluid exiting the heat exchange channels may be at atemperature in the range of about 150° C. to about 400° C., or about200° C. to about 350° C. The residence time of the heat exchange fluidin the heat exchange channels may range from about 1 to about 2000 ms,or about 10 to about 500 ms. The pressure drop for the heat exchangefluid as it flows through the heat exchange channels may range up toabout 10 atm/m, or from about 1 to about 10 atm/m, or from about 3 toabout 7 atm/m, or about 5 atm/m. The heat exchange fluid may be in theform of a vapor, a liquid, or a mixture of vapor and liquid. TheReynolds Number for the flow of the heat exchange fluid in heat exchangechannels may be from about 10 to about 4000, or about 100 to about 2000.

The heat exchange fluid used in the heat exchange channels in themicrochannel reactor core 220 may be any heat exchange fluid suitablefor cooling an exothermic synthesis gas conversion reaction (e.g.,Fischer-Tropsch exothermic reaction). These may include air, steam,liquid water, gaseous nitrogen, other gases including inert gases,carbon monoxide, oils such as mineral oil, and heat exchange fluids suchas Dowtherm A and Therminol which are available from Dow-Union Carbide.

The heat exchange channels used in the microchannel reactor core 220 maycomprise process channels wherein an endothermic process is conducted.These heat exchange process channels may be microchannels. Examples ofendothermic processes that may be conducted in the heat exchangechannels include steam reforming and dehydrogenation reactions. Steamreforming of an alcohol that occurs at a temperature in the range fromabout 200° C. to about 300° C. is an example of an endothermic processthat may be used. The incorporation of a simultaneous endothermicreaction to provide an improved cooling may enable a typical heat fluxof roughly an order of magnitude above convective cooling.

The heat exchange fluid may undergo a partial or full phase change as itflows in the heat exchange channels of the microchannel reactor core220. This phase change may provide additional heat removal from theprocess microchannels beyond that provided by convective cooling. For aliquid heat exchange fluid being vaporized, the additional heat beingtransferred from the Fischer-Tropsch process microchannels may resultfrom the latent heat of vaporization required by the heat exchangefluid. In one embodiment, about 50% by weight of the heat exchange fluidmay be vaporized, or about 35% by weight may be vaporized, or about 20%by weight may be vaporized, or about 10% by weight, or about 5% byweight may be vaporized, or about 2 to about 3% by weight may bevaporized.

The heat flux for heat exchange in the microchannel reactor core 220 maybe in the range from about 0.01 to about 500 watts per square centimeterof surface area of the one or more heat transfer walls of the processmicrochannels (W/cm²) in the microchannel reactor, or in the range fromabout 0.1 to about 250 W/cm², or from about 1 to about 125 W/cm², orfrom about 1 to about 100 W/cm², or from about 1 to about 50 W/cm², orfrom about 1 to about 25 W/cm², or from about 1 to about 10 W/cm². Therange may be from about 0.2 to about 5 W/cm², or about 0.5 to about 3W/cm², or from about 1 to about 2 W/cm².

The control of heat exchange during the synthesis gas conversion process(e.g., Fischer-Tropsch reaction process) may be advantageous forcontrolling selectivity towards the desired product due to the fact thatsuch added cooling may reduce or eliminate the formation of undesiredby-products from undesired parallel reactions with higher activationenergies.

The pressure within each individual heat exchange channel in themicrochannel reactor core 220 may be controlled using passive structures(e.g., obstructions), orifices and/or mechanisms upstream of the heatexchange channels or in the channels. By controlling the pressure withineach heat exchange channel, the temperature within each heat exchangechannel can be controlled. A higher inlet pressure for each heatexchange channel may be used where the passive structures, orificesand/or mechanisms let down the pressure to the desired pressure. Bycontrolling the temperature within each heat exchange channel, thetemperature in the process microchannels can be controlled. Thus, forexample, each process microchannel may be operated at a desiredtemperature by employing a specific pressure in the heat exchangechannel adjacent to or in thermal contact with the process microchannel.This provides the advantage of precisely controlled temperatures foreach process microchannel. The use of precisely controlled temperaturesfor each process microchannel provides the advantage of a tailoredtemperature profile and an overall reduction in the energy requirementsfor the process.

In a scale up device, for certain applications, it may be required thatthe mass of the process fluid be distributed uniformly among themicrochannels. Such an application may be when the process fluid isrequired to be heated or cooled down with adjacent heat exchangechannels. The uniform mass flow distribution may be obtained by changingthe cross-sectional area from one parallel microchannel to anothermicrochannel. The uniformity of mass flow distribution may be defined byQuality Index Factor (Q-factor) as indicated below. A Q-factor of 0%means absolute uniform distribution.

$Q = {\frac{{\overset{.}{m}}_{m{ax}} - {\overset{.}{m}}_{min}}{{\overset{.}{m}}_{m{ax}}} \times 100}$A change in the cross-sectional area may result in a difference in shearstress on the wall. In one embodiment, the Q-factor for the microchannelreactor 110 may be less than about 50%, or less than about 20%, or lessthan about 5%, or less than about 1%.

The superficial velocity for fluid flowing in the process microchannelsmay be at least about 0.01 meters per second (m/s), or at least about0.1 m/s, or in the range from about 0.01 to about 100 m/s, or in therange from about 0.01 to about 10 m/s, or in the range from about 0.1 toabout 10 m/s, or in the range from about 1 to about 100 m/s, or in therange from about 1 to about 10 m/s.

The free stream velocity for fluid flowing in the process microchannelsmay be at least about 0.001 m/s, or at least about 0.01 m/s, or in therange from about 0.001 to about 200 m/s, or in the range from about 0.01to about 100 m/s, or in the range from about 0.01 to about 200 m/s.

The conversion of CO from the fresh synthesis gas may be about 70% orhigher, or about 75% or higher, or about 80% or higher, or about 90% orhigher, or about 91% or higher, or about 92% or higher, or from about88% to about 95%, or from about 90% to about 94%, or from about 91% toabout 93%. If a tail gas recycle is used, the one-pass conversion of COfor the CO in the reactant mixture (i.e., fresh synthesis gas plusrecycled tail gas) may be in the range from about 50% to about 90%, orfrom about 65% to about 85%.

The selectivity to methane in the Fischer-Tropsch (FT) product may be inthe range from about 0.01 to about 10%, or about 1% to about 5%, orabout 1% to about 10%, or about 3% to about 9%, or about 4% to about 8%.

The Fischer-Tropsch product may comprise a gaseous product fraction anda liquid product fraction. The gaseous product fraction may includehydrocarbons boiling below about 350° C. at atmospheric pressure (e.g.,tail gases through middle distillates). The liquid product fraction (thecondensate fraction) may include hydrocarbons boiling above about 350°C. (e.g., vacuum gas oil through heavy paraffins).

The Fischer-Tropsch product fraction boiling below about 350° C. may beseparated into a tail gas fraction and a condensate fraction, e.g.,normal paraffins of about 5 to about 20 carbon atoms and higher boilinghydrocarbons, using, for example, a high pressure and/or lowertemperature vapor-liquid separator, or low pressure separators or acombination of separators. The fraction boiling above about 350° C. (thecondensate fraction) may be separated into a wax fraction boiling in therange of about 350° C. to about 650° C. after removing one or morefractions boiling above about 650° C. The wax fraction may containlinear paraffins of about 20 to about 50 carbon atoms with relativelysmall amounts of higher boiling branched paraffins. The separation maybe effected using fractional distillation.

The Fischer-Tropsch product may include methane, wax and other heavyhigh molecular weight products. The product may include olefins such asethylene, normal and iso-paraffins, and combinations thereof. These mayinclude hydrocarbons in the distillate fuel ranges, including the jet ordiesel fuel ranges.

Branching may be advantageous in a number of end-uses, particularly whenincreased octane values and/or decreased pour points are desired. Thedegree of isomerization may be greater than about 1 mole of isoparaffinper mole of n-paraffin, or about 3 moles of isoparaffin per mole ofn-paraffin. When used in a diesel fuel composition, the product maycomprise a hydrocarbon mixture having a cetane number of at least about60.

The Fischer-Tropsch catalyst may comprise cobalt and a support. Thecatalyst may have a Co loading in the range from about 10 to about 60%by weight, or from about 15 to about 60% by weight, or from about 20 toabout 60% by weight, or from about 25 to about 60% by weight, or fromabout 30 to about 60% by weight, or from about 32 to about 60% byweight, or from about 35 to about 60% by weight, or from about 38 toabout 60% by weight, or from about 40 to about 60% by weight, or fromabout 40 to about 55% by weight, or about 40 to about 50% of cobalt.

The Fischer-Tropsch catalyst may further comprise a noble metal. Thenoble support metal may be one or more of Pd, Pt, Rh, Ru, Re, Ir, Au, Agand Os. The noble metal may be one or more of Pd, Pt, Rh, Ru, Ir, Au, Agand Os. The noble metal may be one or more of Pt, Ru and Re. The noblemetal may be Ru. As an alternative, or in addition, the noble metal maybe Pt. The Fischer-Tropsch catalyst may comprise from about 0.01 toabout 30% in total of noble metal(s) (based on the total weight of allnoble metals present as a percentage of the total weight of the catalystprecursor or activated catalyst), or from about 0.05 to about 20% intotal of noble metal(s), or from about 0.1 to about 5% in total of noblemetal(s), or about 0.2% in total of noble metal(s).

The Fischer-Tropsch catalyst may include one or more other metal-basedcomponents as promoters or modifiers. These metal-based components mayalso be present in the catalyst precursor and/or activated catalyst ascarbides, oxides or elemental metals. A suitable metal for the one ormore other metal-based components may be one or more of Zr, Ti, V, Cr,Mn, Ni, Cu, Zn, Nb, Mo, Tc, Cd, Hf, Ta, W, Re, Hg, TI and the 4f-blocklanthanides. Suitable 4f-block lanthanides may be La, Ce, Pr, Nd, Pm,Sm, Eu, Gd, Tb, Dy, Ho, Er, Tm, Yb and/or Lu. The metal for the one ormore other metal-based components may be one or more of Zn, Cu, Mn, Moand W. The metal for the one or more other metal-based components may beone or more of Re and Pt. The catalyst may comprise from about 0.01 toabout 10% in total of other metal(s) (based on the total weight of allthe other metals as a percentage of the total weight of the catalystprecursor or activated catalyst), or from about 0.1 to about 5% in totalof other metals, or about 3% in total of other metals.

The Fischer-Tropsch catalyst may be derived from a catalyst precursorwhich may be activated to produce the Fischer-Tropsch catalyst, forinstance by heating the catalyst precursor in hydrogen and/or ahydrocarbon gas (e.g., methane), or in a hydrogen or hydrocarbon gasdiluted with another gas, such as nitrogen and/or methane, to convert atleast some of the carbides or oxides to elemental metal. In the activecatalyst, the cobalt may optionally be at least partially in the form ofits carbide or oxide.

The Fischer-Tropsch catalyst precursor may be activated using acarboxylic acid as the reducing agent. The carboxylic acid may be chosensuch that it minimizes the fracturing of the catalyst precursor whilststill ultimately producing an effective catalyst. A mixture of two ormore carboxylic acids may be used. The carboxylic acid may be analpha-hydroxy carboxylic acid, such as citric acid, glycolic acid,lactic acid, mandelic acid, or a mixture of two or more thereof.

The Fischer-Tropsch catalyst may include a catalyst support. The supportmay comprise a refractory metal oxide, carbide, carbon, nitride, ormixture of two or more thereof. The support may comprise alumina,zirconia, silica, titania, or a mixture of two or more thereof. Thesurface of the support may be modified by treating it with silica,titania, zirconia, magnesia, chromia, alumina, or a mixture of two ormore thereof. The material used for the support and the material usedfor modifying the support may be different. The support may comprisesilica and the surface of the silica may be treated with an oxiderefractory solid oxide such as titania. The material used to modify thesupport may be used to increase the stability (e.g. by decreasingdeactivation) of the supported catalyst. The catalyst support maycomprise up to about 30% by weight of the oxide (e.g., silica, titania,magnesia, chromia, alumina, or a mixture of two or more thereof) used tomodify the surface of the support, or from about 1% to about 30% byweight, or from about 5% to about 30% by weight, or from about 5% toabout 25% by weight, or from about 10% to about 20% by weight, or fromabout 12% to about 18% by weight. The catalyst support may be in theform of a structured shape, pellets or a powder. The catalyst supportmay be in the form of particulate solids. While not wishing to be boundby theory, it is believed that the surface treatment provided for hereinhelps keep the Co from sintering during operation of the inventiveFischer-Tropsch process.

The deactivation rate of the Fischer-Tropsch catalyst may be such thatit can be used in a Fischer-Tropsch synthesis for more than about 300hours, or more than about 3,000 hours, or more than about 12,000 hours,or more than about 15,000 hours, all before a catalyst rejuvenation orregeneration is required.

The Fischer-Tropsch catalyst may be used for an extended period(e.g. >300 hours) with a deactivation rate of less than about 1.4% perday, or less than about 1.2% per day, or between about 0.1% and about 1%per day, or between about 0.03 and about 0.15% per day.

The synthesis gas conversion catalyst (e.g., the Fischer-Tropschcatalyst) may have any size and geometric configuration that fits withinthe reactor, e.g., the process microchannels. The catalyst may be in theform of particulate solids (e.g., pellets, powder, fibers, and the like)having a median particle diameter of about 1 to about 1000 μm (microns),or about 10 to about 750 μm, or about 25 to about 500 μm. The medianparticle diameter may be in the range from 50 to about 500 μm or about100 to about 500 μm, or about 125 to about 400 μm, or about 170 to about300 μm. In one embodiment, the catalyst may be in the form of a fixedbed of particulate solids.

The catalyst for the methanol and dimethyl ether reactions may compriseany catalyst suitable for synthesizing methanol or dimethyl ether fromsynthesis gas. These may include catalysts comprising copper, zinc andaluminum oxides (e.g., gamma-alumina), and optionally furthercontaining, for example, oxides of one or more rare earth elements(i.e., elements 57-71), zirconium, yttrium, chromium, silver, gallium,vanadium, molybdenum, tungsten or titanium. The ranges of proportionsmay be from about 30 to about 70% by weight as copper, from about 20 toabout 70% by weight as zinc, and up to about 15% by weight as aluminum.Examples of methanol synthesis catalysts that may be used may includethose disclosed in U.S. Pat. Nos. 4,596,782; 5,238,895; 5,254,520;5,384,335; 5,610,202; 5,767,039; 6,114,279; 6,342,538 B1; 6,433,029 B1;and 6,486,219 B1; and U.S. Patent Publication 2002/0177741 A1.

The dimethyl ether catalysts that may be used may include thosedisclosed in U.S. Pat. Nos. 4,011,275; 6,069,180; 6,147,125; 6,248,795;6,638,892; and J. L. Dubois et al., “Conversion of Carbon Dioxide toDimethyl Ether and Methanol Over Hybrid Catalysts,” Chem. Lett., (7)1115-1118 (1992). These patents and publications are incorporated hereinby reference.

The methanol forming catalyst may be used in combination with adehydration catalyst to provide a synthesis-gas-to-dimethylether route.Examples of the dehydration catalyst that may be used include acidicoxides such as alumina, silica-alumina, zeolite, andsilico-alumino-phosphate synthetic molecular sieves. These are disclosedin U.S. 2006/0020155A1 and US 2007/0244000A1, which are incorporatedherein by reference. The methanol forming catalyst and the dehydrationcatalyst may be mixed or combined together in the same reaction zone.Alternatively, the dehydration catalyst may be positioned downstream ofthe methanol forming catalyst, either in the same microchannel reactoror in a separate microchannel reactor.

The catalyst used for the methane forming reactions may comprise anycatalyst suitable for converting synthesis gas to methane. The catalystmay comprise nickel, iron, cobalt, ruthenium, molybdenum, vanadium,titanium, or a mixture of two or more thereof. The catalyst may comprisean oxide of any of the foregoing metals. The catalyst may comprisevanadium and/or molybdenum in the form of free metal, salt, oxide and/orsulfide on a porous, oxidic support comprising titanium dioxide. Thecatalyst may be promoted with one or more salts, hydroxides, oxides orsulfides of one or more metals belonging to Groups IA, IIA or IIIB ofthe Periodic Table. The catalyst may comprise vanadium sulfide promotedwith ceruim sulfide on a porous support comprising titanium dioxide. Thecatalysts that may be used are described in U.S. Pat. No. 4,540,714,which is incorporated herein by reference.

The catalyst may be in the form of a fixed bed of particulate solids (asshown in FIG. 8 ). Referring to FIG. 8 , the catalyst 400, which is inthe form of a bed of particulate solids, is contained in processmicrochannel 402. Reactants enter the fixed bed as indicated by arrow404, undergo reaction, and product flows out of the fixed bed asindicated by arrow 406.

The catalyst may be supported on a catalyst support structure such as afoam, felt, wad or a combination thereof. The catalyst support structuremay comprise a fin assembly or corrugated inserts suitable for insertioninto slots in the microchannel reactor.

The term “foam” is used herein to refer to a structure with continuouswalls defining pores throughout the structure. The term “felt” is usedherein to refer to a structure of fibers with interstitial spacestherebetween. The term “wad” is used herein to refer to a structure oftangled strands, like steel wool. The catalyst may be supported on ahoneycomb structure. The catalyst may be supported on a flow-by supportstructure such as a felt with an adjacent gap, a foam with an adjacentgap, a fin structure with gaps, a washcoat on any inserted substrate, ora gauze that is parallel to the flow direction with a corresponding gapfor flow.

An example of a flow-by structure is illustrated in FIG. 9 . In FIG. 9 ,the catalyst 410 is contained within process microchannel 412. An openpassage way 414 permits the flow of fluid through the processmicrochannel 412 as indicated by arrows 416 and 418. The reactantscontact the catalyst and undergo reaction to form product.

The catalyst may be supported on a flow-through support structure suchas a foam, wad, pellet, powder, or gauze. An example of a flow-throughstructure is illustrated in FIG. 10 . In FIG. 10 , the flow-throughcatalyst 420 is contained within process microchannel 422, the reactantsflow through the catalyst 420 as indicated by arrow 424, and undergoreaction to form the product 426.

The support structure for a flow-through catalyst may be formed from amaterial comprising silica gel, foamed copper, sintered stainless steelfiber, steel wool, alumina, or a combination of two or more thereof. Thesupport structure may be made of a heat conducting material, such as ametal, to enhance the transfer of heat to or from the catalyst.

The catalyst may be supported on a fin assembly comprising one or morefins positioned within the process microchannels. Examples areillustrated in FIGS. 11-13 . Referring to FIG. 11 , fin assembly 480includes fins 481 which are mounted on fin support 483 which overliesbase wall 484 of process microchannel 485. The fins 481 project from thefin support 483 into the interior of the process microchannel 485. Thefins 481 may extend to and contact the interior surface of upper wall486 of process microchannel 485. Fin channels 487 between the fins 481provide passage ways for reactant and product to flow through theprocess microchannel 485 parallel to its length. Each of the fins 481has an exterior surface on each of its sides. The exterior surfaceprovides a support base for the catalyst. The reactants may flow throughthe fin channels 487, contact the catalyst supported on the exteriorsurface of the fins 481, and react to form product. The fin assembly 480a illustrated in FIG. 12 is similar to the fin assembly 480 illustratedin FIG. 11 except that the fins 481 a do not extend all the way to theinterior surface of the upper wall 486 of the microchannel 485. The finassembly 480 b illustrated in FIG. 13 is similar to the fin assembly 480illustrated in FIG. 11 except that the fins 481 b in the fin assembly280 b have cross sectional shapes in the form of trapezoids. Each of thefins may have a height ranging from about 0.02 mm up to the height ofthe process microchannel 485, or from about 0.02 to about 10 mm, or fromabout 0.02 to about 5 mm, or from about 0.02 to about 2 mm. The width ofeach fin may range from about 0.02 to about 5 mm, or from about 0.02 toabout 2 mm, or about 0.02 to about 1 mm. The length of each fin may beof any length up to the length of the process microchannel 485, or up toabout 3 m, or about 0.03 to about 3 m, or about 0.03 to about 2.5 m, orabout 0.03 to about 2 m. The gap between each of the fins may be of anyvalue and may range from about 0.02 to about 5 mm, or from about 0.02 toabout 2 mm, or from about 0.02 to about 1 mm. The number of fins in theprocess microchannel 485 may range from about 1 to about 50 fins percentimeter of width of the process microchannel 285, or from about 1 toabout 30 fins per centimeter, or from about 1 to about 10 fins percentimeter, or from about 1 to about 5 fins per centimeter, or fromabout 1 to about 3 fins per centimeter. Each of the fins may have across-section in the form of a rectangle or square as illustrated inFIG. 11 or 12 , or a trapezoid as illustrated in FIG. 13 . When viewedalong its length, each fin may be straight, tapered or have a serpentineconfiguration. The fin assembly may be made of any material thatprovides sufficient strength, dimensional stability and heat transfercharacteristics to permit operation for which the process microchannelis intended. These materials include: steel (e.g., stainless steel,carbon steel, and the like); aluminum; titanium; nickel; platinum;rhodium; copper; chromium; alloys of any of the foregoing metals; monel;inconel; brass; polymers (e.g., thermoset resins); ceramics; glass;quartz; silicon; or a combination of two or more thereof. The finassembly may be made of an Al₂O₃ or a Cr₂O₃ forming material wherein alayer of Al₂O₃ or a Cr₂O₃ forms on the surface of the fin assembly whenthe fin assembly is heat treated in air. The fin assembly may be made ofan alloy comprising Fe, Cr, Al and Y, or an alloy comprising Ni, Cr andFe.

The catalyst may be supported on one or more corrugated insertspositioned in slots within the microchannel reactor. This is illustratedin FIG. 14 wherein microchannel reactor 500 includes corrugated inserts502 inserted in slots 504. The slots 504 may comprise microchannels, andhave the dimensions indicated above as being microchannels.Alternatively, the slots 504 may have dimensions that would make themlarger than microchannels. The process microchannels of the microchannelreactor may comprise the slots 504, or may be positioned within thecorrugated inserts 502 and/or formed by openings between the interiorsidewalls of the slots 504 and the inserts 502. Each of the corrugatedinserts 502 may have a height ranging from about 0.02 mm up to theheight of the slot 504, or from about 0.02 to about 10 mm, or from about0.02 to about 5 mm, or from about 0.02 to about 2 mm. Each of thecorrugated inserts 502 may have a width ranging from about 0.02 mm up tothe width of the slot 504, or from about 0.02 to about 10 mm, or fromabout 0.02 to about 5 mm, or from about 0.02 to about 2 mm. The lengthof each corrugated insert may be of any length up to the length of theslot 504, or up to about 3 m, or about 0.03 to about 3 m, or about 0.03to about 2 m, or about 0.03 to about 1 m. The corrugated inserts 502 maybe made of any material that provides sufficient strength, dimensionalstability and heat transfer characteristics to permit operation forwhich the microchannel reactor is intended. These materials include:steel (e.g., stainless steel, carbon steel, and the like); aluminum;titanium; nickel; platinum; rhodium; copper; chromium; alloys of any ofthe foregoing metals; monel; inconel; brass; polymers (e.g., thermosetresins); ceramics; glass; quartz; silicon; or a combination of two ormore thereof. The corrugated inserts 502 may be made of an alloy thatforms a layer of Al₂O₃ or Cr₂O₃ on the surface of the inserts when heattreated in air. The corrugated inserts 502 may be made of an alloycomprising Fe, Cr, Al and Y, or an alloy comprising Ni, Cr and Fe.

The catalyst may be directly washcoated or grown from solution on theinterior walls of the process microchannels and/or on one or more of theabove-described catalyst support structures. The catalyst may be in theform of a single piece of porous contiguous material, or a plurality ofpieces in physical contact. The catalyst may comprise a contiguousmaterial and have a contiguous porosity such that molecules can diffusethrough the catalyst. In this embodiment, the fluids may flow throughthe catalyst rather than around it. The cross-sectional area of thecatalyst may occupy from about 1 to about 99%, or about 10 to about 95%of the cross-sectional area of the process microchannels.

The catalyst may comprise a support, an interfacial layer on thesupport, and a catalyst material on or mixed with the interfacial layer.The support may comprise one or more of the above-described foams,felts, wads, fin structures, or corrugated inserts. The interfaciallayer may be solution deposited on the support or it may be deposited bychemical vapor deposition or physical vapor deposition. The catalyst maycomprise the support, a buffer layer, an interfacial layer, and thecatalyst material. The support may be porous. Any of the foregoinglayers may be continuous or discontinuous as in the form of spots ordots, or in the form of a layer with gaps or holes. The support may havea porosity of at least about 5% as measured by mercury porosimetry andan average pore size (sum of pore diameters divided by number of pores)of about 1 to about 2000 microns, or from about 1 to about 1000 microns.The support may be a porous ceramic or a metal foam. Other supports thatmay be used may include carbides, nitrides, and composite materials. Thesupport may have a porosity of about 30% to about 99%, or about 60% toabout 98%. The support may be in the form of a foam, felt, wad, or acombination thereof. The open cells of the metal foam may range fromabout 20 pores per inch (ppi) to about 3000 ppi, and in one embodimentabout 20 to about 1000 ppi, and in one embodiment about 40 to about 120ppi. The term “ppi” refers to the largest number of pores per inch (inisotropic materials the direction of the measurement is irrelevant;however, in anisotropic materials, the measurement is done in thedirection that maximizes pore number).

The buffer layer, when present, may have a different composition and/ordensity than both the support and the interfacial layers, and in oneembodiment has a coefficient of thermal expansion that is intermediatethe thermal expansion coefficients of the porous support and theinterfacial layer. The buffer layer may be a metal oxide or metalcarbide. The buffer layer may comprise Al₂O₃, TiO₂, SiO₂, ZrO₂, orcombination thereof. The Al₂O₃ may be α-Al₂O₃, γ-Al₂O₃ or a combinationthereof. The buffer layer may comprise an oxide layer (e.g. Al₂O₃ orCr₂O₃) formed by heat treating the support in air. The buffer layer maybe formed of two or more compositionally different sublayers. Forexample, when the porous support is metal, for example a stainless steelfoam, a buffer layer formed of two compositionally different sub-layersmay be used. The first sublayer (in contact with the porous support) maybe TiO₂. The second sublayer may be α-Al₂O₃ which is placed upon theTiO₂. In one embodiment, the α-Al₂O₃ sublayer is a dense layer thatprovides protection of the underlying metal surface. A less dense, highsurface area interfacial layer such as alumina may then be deposited assupport for a catalytically active layer.

The support may have a thermal coefficient of expansion different fromthat of the interfacial layer. In such a case a buffer layer may beneeded to transition between the two coefficients of thermal expansion.The thermal expansion coefficient of the buffer layer can be tailored bycontrolling its composition to obtain an expansion coefficient that iscompatible with the expansion coefficients of the porous support andinterfacial layers. The buffer layer should be free of openings and pinholes to provide superior protection of the underlying support. Thebuffer layer may be nonporous. The buffer layer may have a thicknessthat is less than one half of the average pore size of the poroussupport. The buffer layer may have a thickness of about 0.05 to about 10μm, or about 0.05 to about 5 μm.

In an embodiment adequate adhesion and chemical stability may beobtained without a buffer layer. In this embodiment the buffer layer maybe omitted.

The interfacial layer may comprise nitrides, carbides, sulfides,halides, metal oxides, carbon, or a combination thereof. The interfaciallayer provides high surface area and/or provides a desirablecatalyst-support interaction for supported catalysts. The interfaciallayer may be comprised of any material that is conventionally used as acatalyst support. The interfacial layer may comprise a metal oxide.Examples of metal oxides that may be used include α-Al₂O₃, SiO₂, ZrO₂,TiO₂, tungsten oxide, magnesium oxide, vanadium oxide, chromium oxide,manganese oxide, iron oxide, nickel oxide, cobalt oxide, copper oxide,zinc oxide, molybdenum oxide, tin oxide, calcium oxide, aluminum oxide,lanthanum series oxide(s), zeolite(s) and combinations thereof. Theinterfacial layer may serve as a catalytically active layer without anyfurther catalytically active material deposited thereon. The interfaciallayer may be used in combination with a catalytically active layer. Thecatalyst may be mixed with the interfacial layer. The interfacial layermay also be formed of two or more compositionally different sublayers.The interfacial layer may have a thickness that is less than one half ofthe average pore size of the porous support. The interfacial layerthickness may range from about 0.5 to about 100 μm, and in oneembodiment from about 1 to about 50 microns. The interfacial layer maybe either crystalline or amorphous. The interfacial layer may have a BETsurface area of at least about 1 m²/g.

The catalyst may be deposited on the interfacial layer. Alternatively,the catalyst may be simultaneously deposited with the interfacial layer.The catalyst layer may be intimately dispersed on the interfacial layer.That the catalyst layer is “dispersed on” or “deposited on” theinterfacial layer includes the conventional understanding thatmicroscopic catalyst particles are dispersed: on the support layer (i.e., interfacial layer) surface, in crevices in the support layer, and inopen pores in the support layer.

The catalyst may be in the form of a bed of particulates. The catalyticbed of particulate solids may be graded in composition or graded with athermally conductive inert material. The thermally conductive inertmaterial may be interspersed with the active catalyst. Examples ofthermally conductive inert materials that may be used include diamondpowder, silicon carbide, aluminum, alumina, copper, graphite, and thelike. The catalyst bed fraction may range from about 100% by weightactive catalyst to less than about 50% by weight active catalyst. Thecatalyst bed fraction may range from about 10% to about 90% by weightactive catalyst, and in one embodiment from about 25% to about 75% byweight. In an alternate embodiment the thermally conductive inertmaterial may be deployed at the center of the catalyst or within thecatalyst particles. The active catalyst may be deposited on the outside,inside or intermittent within a composite structure that includes thethermally conductive inert. The resultant catalyst composite structuremay have an effective thermal conductivity when placed in a processmicrochannel or combustion channel that is at least about 0.3 W/m/K, andin one embodiment at least about 1 W/m/K, and in one embodiment at leastabout 2 W/m/K.

The catalyst bed may be graded only locally within the processmicrochannel. For example, a process microchannel may contain a catalystbed with a first reaction zone and a second reaction zone. The top orbottom (or front or back) of the catalyst bed may be graded incomposition whereby a more or less active catalyst is employed in all orpart of the first or second reaction zone. The composition that isreduced in one reaction zone may generate less heat per unit volume andthus reduce the hot spot and potential for the production of undesirableby-products, such as methane in a Fischer-Tropsch reaction. The catalystmay be graded with an inert material in the first and/or second reactionzone, in full or in part. The first reaction zone may contain a firstcomposition of catalyst or inert material, while the second reactionzone may contain a second composition of catalyst or inert material.

Different particle sizes may be used in different axial regions of theprocess microchannels to provide for graded catalyst beds. For example,very small particles may be used in a first reaction zone while largerparticles may be used in a second reaction zone. The average particlediameters may be less than half the height or gap of the processmicrochannels. The very small particles may be less than one-fourth ofthe process microchannel height or gap. Larger particles may cause lowerpressure drops per unit length of the process microchannels and may alsoreduce the catalyst effectiveness. The effective thermal conductivity ofa catalyst bed may be lower for larger size particles. Smaller particlesmay be used in regions where improved heat transfer is sought throughoutthe catalyst bed or alternatively larger particles may be used to reducethe local rate of heat generation.

Relatively short contact times, high selectivity to the desired productand relatively low rates of deactivation of the catalyst may be achievedby limiting the diffusion path required for the catalyst. This may beachieved when the catalyst is in the form of a thin layer on anengineered support such as a metallic foam or on the wall of the processmicrochannel. This may allow for increased space velocities. The thinlayer of catalyst may be produced using chemical vapor deposition. Thisthin layer may have a thickness in the range up to about 1 micron, andin one embodiment in the range from about 0.1 to about 1 micron, and inone embodiment in the range from about 0.1 to about 0.5 micron, and inone embodiment about 0.25 micron. These thin layers may reduce the timethe reactants are within the active catalyst structure by reducing thediffusional path. This may decrease the time the reactants spend in theactive portion of the catalyst. The result may be increased selectivityto the product and reduced unwanted by-products. An advantage of thismode of catalyst deployment may be that, unlike conventional catalystsin which the active portion of the catalyst may be bound up in an inertlow thermal conductivity binder, the active catalyst film may be inintimate contact with either an engineered structure or a wall of theprocess microchannel. This may leverage high heat transfer ratesattainable in the microchannel reactor and allow for close control oftemperature. This may result in the ability to operate at increasedtemperature (faster kinetics) without promoting the formation ofundesired by-products, thus producing higher productivity and yield andprolonging catalyst life.

The configuration of the microchannel reactor may be tailored to matchthe reaction kinetics. Near the entrance or top of a first reaction zoneof a process microchannel, the microchannel height or gap may be smallerthan in a second reaction zone near the exit or bottom of the processmicrochannel. Alternatively, the reaction zones may be smaller than halfthe process microchannel length. For example, a first processmicrochannel height or gap may be used for the first 25%, 50%, 75%, or90% of the length of the process microchannel for a first reaction zone,while a larger second height or gap may be used in a second reactionzone downstream from the first reaction zone. This configuration may besuitable for conducting Fischer-Tropsch reactions. Other gradations inthe process microchannel height or gap may be used. For example, a firstheight or gap may be used near the entrance of the microchannel toprovide a first reaction zone, a second height or gap downstream fromthe first reaction zone may be used to provide a second reaction zone,and a third height or gap may be used to provide a third reaction zonenear the exit of the microchannel. The first and third heights or gapsmay be the same or different. The first and third heights or gaps may belarger or smaller than the second height or gap. The third height or gapmay be smaller or larger than the second height or gap. The secondheight or gap may be larger or smaller than the third height or gap.

The catalyst may be rejuvenated or regenerated by flowing a rejuvinatingor regenerating fluid through the process microchannels in contact withthe catalyst. The rejuvenating or regenerating fluid may comprisehydrogen or a diluted hydrogen stream. The diluent may comprisenitrogen, argon, helium, methane, carbon dioxide, steam, or a mixture oftwo or more thereof. The temperature of the rejuvenating or regeneratingfluid may be from about 50 to about 400° C., and in one embodiment about200 to about 350° C. The pressure within the channels during thisrejuvenation or regeneration step may range from about 1 to about 40atmospheres, and in one embodiment about 1 to about 20 atmospheres, andin one embodiment about 1 to about 5 atmospheres. To complete theregeneration of the catalyst, after the rejuvenating or regeneratingfluid flows in contact with the catalyst, the catalyst may be oxidizedand then reduced. The residence time for the rejuvenating orregenerating fluid in the channels may range from about 0.01 to about1000 seconds, and in one embodiment about 0.1 second to about 100seconds. The microchannel reactor may comprise a reactant header, aproduct footer and a plurality of process microchannels connecting tothe header and footer. The rejuvenating or regenerating fluid may flowfrom the header through the process microchannels to the footer.Alternatively, the rejuvenating or the regenerating fluid may flow fromthe footer through the process microchannels to the header.

The catalyst may be rejuvenated by removing wax and other hydrocarbonsfrom the catalyst. The catalyst may be regenerated by removing wax andother hydrocarbons from the catalyst (typically by stripping with H₂),oxidizing the catalyst with air or other O₂ containing gas at anelevated temperature, re-reducing the catalyst, and then activating thecatalyst.

A synthesis gas conversion production run (e.g., Fischer-Tropschproduction run) will typically be conducted over an extended period oftime, for example, at least about 300 hours, or at least about 3000hours, or at least about 12,000 hours, or at least about 15,000 hours.The reactivity of the catalyst will typically decline over time and tomake up for this decline the temperature of the reaction is usuallyincreased in order to maintain a constant level of production. However,at some point in time the production run will be stopped, eitherpurposely or accidentally, and the problem addressed with this inventionrelates to providing a rapid restart of the production run. With thisinvention, the production run may be restarted within a period of, forexample, up to about 3 hours, or up to about 2 hours, or up to about 1hour, or up to about 0.1 hour from the time the flow of synthesis gasinto the reactor is restarted.

In an embodiment, this invention relates to a synthesis gas conversionprocess (e.g., Fischer-Tropsch process) which comprises flowingsynthesis gas into a reactor in contact with a Fischer-Tropsch catalystat a reaction temperature and pressure to produce a Fischer-Tropschproduct. The flow of synthesis gas entering the reactor may be stoppedfor up to about 48 hours, or up to about 36 hours, or up to about 24hours, or up to about 12 hours, or up to about 3 hours, or up to about 1hour. This may be referred to as a bottled period. The Fischer-Tropschprocess may then be restarted by a method comprising: restoring thepressure within the reactor at the reaction pressure, maintaining thetemperature within the reactor between the coolant temperature and thereaction temperature with the coolant temperature set close (e.g.,within about 15° C., or about 10° C., or about 5° C., or about 1° C.below) to the original reaction temperature; and restarting the flow ofsynthesis gas into the reactor, and the flow of effluent out of thereactor. In this embodiment the return to full production may be almostinstantaneous. The reaction temperature may be controlled with a coolantflowing in a heat exchanger in thermal contact with the reactor. Thebottling may be followed by purging with hydrogen, nitrogen ordesulfurized natural gas and reintroduction of synthesis gas into thereactor.

In an embodiment, the flow of synthesis gas entering the reactor isstopped. The Fischer-Tropsch process may be restarted by a methodcomprising: restoring the pressure within the reactor at the reactionpressure, and restarting the flow of synthesis gas into the reactor. Thereactor temperature may then be ramped up to the original operatingtemperature, prior to the stop, within a period of time in the rangefrom 0.1 to about 24 hours, or about 1 to about 3 hours, by heating thereactor at a rate of up to about 5° C. per hour, or up to about 10° C.per hour, or up to about 15° C. per hour, or up to about 30° C. perhour, or up to about 60° C. per hour. The reaction temperature may becontrolled with a heat exchange fluid flowing in a heat exchanger inthermal contact with the reactor

In an embodiment, the flow of synthesis gas entering the reactor isstopped. The Fischer-Tropsch process may be restarted by a methodcomprising: flowing hydrogen into the reactor to purge the reactor ofreactants and effluent, holding the reactor in a hydrogen environmentfor a period of time of up to about 48 hours, or up to about 36 hours,or up to about 24 hours, or up to about 12 hours, and then restartingthe flow of synthesis gas into the reactor. The reaction temperature maybe in the range from about 150° C. to about 300° C., and during the stepof flowing hydrogen into the reactor the temperature within the reactormay be increased to a temperature above the reaction temperature, forexample, up to about 350° C., or up to about 400° C. This may rejuvenatethe catalyst. The reaction temperature may be controlled with a heatexchange fluid flowing in a heat exchanger in thermal contact with thereactor.

In an embodiment, the flow of synthesis gas into the reactor is stopped.The Fischer-Tropsch process may be restarted by a method comprising:flowing desulfurized natural gas into the reactor to purge the reactorof reactants and effluent, holding the reactor in a natural gasenvironment for a period of time of up to about 48 hours, or up to about36 hours, or up to about 24 hours, or up to about 12 hours, or up toabout 6 hours, or up to about 3 hours, or up to about 1 hour and thenrestarting the flow of synthesis gas into the reactor and the flow ofeffluent out of the reactor.

In an embodiment, the flow of synthesis gas into the reactor is stopped.The Fischer-Tropsch process may be restarted by a method comprising:maintaining the reactor at the pre-stop operating temperature for aperiod of time of up to about 48 hours, or up to about 36 hours, or upto about 24 hours, or up to about 12 hours, or up to about 6 hours, orup to about 3 hours, or up to about 1 hour; restoring the pressure tothe operating pressure (if needed); and flowing hydrogen, nitrogen, ordesulfurized natural gas into the reactor to purge the reactor ofreactants and effluent; holding the reactor in a purge gas environmentfor a period of time of up to about 48 hours, or up to about 36 hours,or up to about 24 hours, or up to about 12 hours, or up to about 6hours, or up to about 3 hours, or up to about 1 hour; and thenrestarting the flow of synthesis gas into the reactor and the flow ofeffluent out of the reactor. In an embodiment, the purge gas isintroduced into the reactor as close to the catalyst bed as possible tominimize the additional fresh reactant supply in contact with thecatalyst.

In an embodiment, the flow of synthesis gas entering the reactor isstopped. The Fischer-Tropsch process may be restarted by a methodcomprising: maintaining the reactor at the pre-stop operatingtemperature for a period of up to about 48 hours, or up to about 36hours, or up to about 24 hours, or up to about 12 hours, or up to about6 hours, or up to about 3 hours, or up to about 1 hour; restoring thepressure to the operating pressure (if needed); and flowing hydrogen,nitrogen, or desulfurized natural gas into the reactor to purge thereactor of reactants and effluent; holding the reactor in a purge gasenvironment for a period of time of up to about 48 hours, or up to about36 hours, or up to about 24 hours, or up to about 12 hours, or up toabout 6 hours, or up to about 3 hours, or up to about 1 hour; coolingthe reactor to a temperature lower than the operating temperature; andthen restarting the flow of synthesis gas into the reactor and the flowof effluent out of the reactor. The reactor temperature may then beramped up to the operating temperature used prior to the stop within aperiod of time of up to about 24 hours, or up to about 12 hours, or upto about 6 hours, or up to about 3 hours, or up to about 1 hour, or upto about 0.1 hour, by heating the reactor at a rate of up to about 5° C.per hour, or up to about 10° C. per hour, or up to about 15° C. perhour, or up to about 30° C. per hour, or up to about 60° C. per hour.Optionally, purging at a pressure less than operating pressure may beconducted, with the operating pressure restored when introducingsynthesis gas.

In an embodiment, the flow of synthesis gas into the reactor is stopped.The Fischer-Tropsch process may be restarted by a method comprising:maintaining the reactor at the pre-stop operating temperature for aperiod of time of up to about 48 hours, or up to about 36 hours, or upto about 24 hours, or up to about 12 hours, or up to about 6 hours, orup to about 3 hours, or up to about 1 hour; depressurizing the reactorto a pressure (lower than the operating pressure) of up to about 20atmospheres, or up to about 10 atmospheres, or up to about 5atmospheres; and flowing hydrogen, nitrogen, or desulfurized natural gasinto the reactor to purge the reactor of reactants and effluent; holdingthe reactor in a purge gas environment for a period of time of up toabout 48 hours, or up to about 36 hours, or up to about 24 hours, or upto about 12 hours, or up to about 6 hours, or up to about 3 hours, or upto about 1 hour; restarting the flow of synthesis gas into the reactorand effluent out of the reactor; and pressurizing the reactor to thetarget operating pressure. If the reactor temperature is reduced to thatbelow the pre-stop operating temperature, the reactor temperature maythen be ramped up to the operating temperature prior to the stop withina period of time of up to about 48 hours, or up to about 36 hours, or upto about 24 hours, or up to about 12 hours, or up to about 6 hours, orup to about 3 hours, or up to about 1 hour, or up to about 0.1 hour, byheating at a rate of up to about 15° C. per hour, or up to about 30° C.per hour, or up to about 60° C. per hour.

In an embodiment, the flow of synthesis gas entering the reactor isstopped. The Fischer-Tropsch process may be restarted by a methodcomprising: restarting the flow of synthesis gas into the reactor andthe flow of effluent out of the reactor, the temperature of thesynthesis gas flowing into the reactor being within about 10° C. of thereaction temperature in the reactor, or within about 5° C. of thereaction temperature. The reaction temperature may be controlled with aheat exchange fluid flowing in a heat exchanger in thermal contact withthe reactor, and during the step of restarting the flow of synthesis gasinto the reactor and the flow of effluent out of the reactor, thetemperature of the heat exchange fluid in the heat exchanger may be lessthan about 10° C., or less than about 5° C., of the reaction temperaturein the reactor. In this embodiment, the Fischer-Tropsch catalyst maycomprise a wet catalyst.

In an embodiment, the flow of synthesis gas into the reactor is stopped.The Fischer-Tropsch process may be restarted by a method comprising:rejuvenating the catalyst in a hydrogen environment, or regenerating thecatalyst by de-waxing the catalyst followed by oxidation and reductionof the catalyst; and restarting the flow of synthesis gas into thereactor and the flow of effluent out of the reactor, the temperature ofthe synthesis gas flowing into the reactor being at the reactiontemperature. The reaction temperature in the reactor may be controlledwith a heat exchange fluid flowing in a heat exchanger in thermalcontact with the reactor, and during the restarting of the flow ofsynthesis gas into the reactor, the temperature of the heat exchangefluid in the heat exchanger may be lower than the reaction temperaturein the reactor, for example, up to about 10° C. lower, or up to about 5°C. lower, than the reaction temperature in the reactor.

In an embodiment, the flow of synthesis gas into the reactor is stopped.The catalyst in the reactor is a wet catalyst. The Fischer-Tropschprocess may be restarted by a method comprising: flowing hydrogen at atemperature of up to about 400° C. to regenerate the catalyst, followedby an oxidation treatment involving flowing an oxygen containing gasmixture, such as air (21% O₂), into the reactor in contact with theFischer-Tropsch catalyst at a temperature from about 70° C. to about350° C., or from about 250° C. to about 300° C. for a period of time inthe range from about 1 to about 12 hours. This may be followed by acatalyst reduction treatment involving flowing hydrogen at a temperatureof up to about 400° C. in contact with the catalyst to regenerate theFischer-Tropsch catalyst. The flow of synthesis gas into the reactor incontact with the regenerated catalyst and the flow of effluent out ofthe reactor may then be started. The reaction temperature in the reactormay be controlled with a heat exchange fluid flowing in a heat exchangerin thermal contact with the reactor, and during the restarting of theflow of synthesis gas into the reactor, the temperature of the heatexchange fluid in the heat exchanger may be lower than the reactiontemperature in the reactor, for example, up to about 10° C. lower, or upto about 5° C. lower, than the reaction temperature in the reactor.

In an embodiment, the flow of synthesis gas entering the reactor isstopped. The Fischer-Tropsch process may be restarted by a methodcomprising: selectively infusing hydrogen into the reactor without thepurging of the synthesis gas in contact with the catalyst and/orselectively defusing (removing) water vapor from the process gas mixturein the reactor and then restarting the flow of synthesis gas into thereactor and the flow of effluent out of the reactor and adjusting theprocess pressures and temperatures to the pre-stoppage levels.Conditions should be avoided which would cause the condensation ofliquid water in the reactor at the reactor temperature and pressure. Forexample, the partial pressure of water in the mixture should not exceedthe saturation vapor pressure of water (at reactor temperature andpressure) even if all CO present in the mixture is converted, forming amole of water for every mole of CO initially present in the feed.

In an embodiment, the flow of synthesis gas entering the reactor isstopped and the flow of effluent exiting the reactor is stopped. Theprocess (e.g., Fischer-Tropsch process) may be restarted by a methodcomprising: flowing hydrogen into the reactor to purge the reactor ofreactants and effluent at a volumetric flow rate which is equal to orhigher than the volumetric flow rate of synthesis gas prior to thesynthesis gas flow stoppage. Maintaining such a flow rate until thereactor has been purged helps prevent high conversion products fromforming which could negatively affect apparent catalyst activity. Theflow of synthesis gas into the reactor and the flow of effluent out ofthe reactor is then restarted.

In an embodiment, the flow of synthesis gas entering the reactor isstopped and the flow of effluent exiting the reactor is stopped. Theprocess (e.g., Fischer-Tropsch process) may be restarted by a methodcomprising: flowing nitrogen gas into the reactor to purge the reactorof reactants and effluent at a volumetric flow rate which is equal to orhigher than the volumetric flow rate of synthesis gas prior to thesynthesis gas flow stoppage. The flow of synthesis gas into the reactorand the flow of effluent out of the reactor is then restarted.

In an embodiment, the flow of synthesis gas entering the reactor isstopped and the flow of effluent exiting the reactor is stopped. Theprocess (e.g., Fischer-Tropsch process) may be restarted by a methodcomprising: flowing natural gas (e.g., desulfurized natural gas) intothe reactor to purge the reactor of reactants and effluent at avolumetric flow rate which is equal to or higher than the volumetricflow rate of synthesis gas prior to the synthesis gas flow stoppage. Theflow of synthesis gas into the reactor and the flow of effluent out ofthe reactor is then restarted.

In an embodiment, the flow of synthesis gas entering the reactor isstopped. This is followed by flowing hydrogen at a temperature lowerthan the temperature of the reactor into the reactor to purge thereactor of reactants and effluent. The flow of synthesis gas into thereactor is then restarted.

In an embodiment, the flow of synthesis gas entering the reactor isstopped. This is followed by flowing nitrogen gas at a temperature lowerthan the temperature of the reactor into the reactor to purge thereactor of reactants and effluent. The flow of synthesis gas into thereactor is then restarted.

In an embodiment, the flow of synthesis gas entering the reactor isstopped. This is followed by flowing natural gas (e.g., desulfurizednatural gas) at a temperature lower than the temperature of the reactorinto the reactor to purge the reactor of reactants and effluent. Theflow of synthesis gas into the reactor is then restarted.

In an embodiment, the flow of synthesis gas entering the reactor isstopped and the flow of effluent exiting the reactor is stopped. This isfollowed by flowing hydrogen gas into the reactor to purge the reactorof reactants and effluent where the volume of purge gas used is equal toor higher than the volume of synthesis gas in the system between thelocations for stopping the flow of synthesis gas into the reactor andthe stopping the flow of effluent out of the reactor. This is followedby restarting the flow of synthesis gas into the reactor and the flow ofeffluent out of the reactor.

In an embodiment, the flow of synthesis gas entering the reactor isstopped and the flow of effluent exiting the reactor is stopped. This isfollowed by flowing nitrogen gas into the reactor to purge the reactorof reactants and effluent where the volume of purge gas used is equal toor higher than the volume of synthesis gas in the system between thelocations for stopping the flow of synthesis gas into the reactor andstopping the flow of the effluent out of the reactor. This is followedby restarting the flow of synthesis gas into the reactor and the flow ofeffluent out of the reactor.

In an embodiment, the flow of synthesis gas entering the reactor isstopped and the flow of effluent exiting the reactor is stopped. This isfollowed by flowing natural gas (e.g., desulfurized natural gas) intothe reactor to purge the reactor of reactants and effluent where thevolume of purge gas used is equal to or higher than the volume ofsynthesis gas in the system between the locations for stopping the flowof synthesis gas into the reactor and stopping the flow of the effluentout of the reactor. This is followed by restarting the flow of synthesisgas into the reactor and the flow of effluent out of the reactor.

This invention relates to a method of operating a Fischer-Tropschprocess wherein the process is conducted in a plant comprising aplurality of reaction trains. Each reaction train may comprise at leastone Fischer-Tropsch reactor containing a Fischer-Tropsch catalyst, thereaction trains being connected to a common reactant feed streamcomprising fresh synthesis gas, and the flow of synthesis gas into theFischer-Tropsch reactor and the flow of effluent out of theFischer-Tropsch reactor of one or more of the reaction trains isstopped, while the flow of synthesis gas into and the flow of effluentout of the remainder of the reaction trains in the plant is continued.The method comprises: (A) flowing the reactant feed stream at an overallprocess flow rate to the plurality of reaction trains in the plant; (B)dividing the reactant feed stream into a plurality of reactantsubstreams; (C) flowing each reactant substream through a separatereaction train to convert the reactants in the reactant substream to aFischer-Tropsch product; (D) stopping the flow of a reactant substreamto one or more of the reaction trains; and (E) continuing to flow thereactant feed stream to the remainder of reaction trains in the plant atthe same overall process flow rate. During step (C) a mixture of freshsynthesis gas and a recycled tail gas may flow into the Fischer-Tropschreactor of each reaction train. During step (E) the flow of recycledtail gas into the Fischer-Tropsch reactor of the one or more of theremainder of the reaction trains in the plant may be stopped or reducedin order to allow for maintaining the overall flow of fresh synthesisgas into the plant at a constant level.

This invention relates to a method of operating a Fischer-Tropschprocess wherein the process is conducted in a plant comprising aplurality of reaction trains, each reaction train comprising at leastone Fischer-Tropsch reactor containing a Fischer-Tropsch catalyst, thereaction trains being connected to a reactant feed stream comprisingfresh synthesis gas, and the flow of synthesis gas into and the flow ofeffluent out of the Fischer-Tropsch reactor of one or more of thereaction trains is stopped, while the flow of synthesis gas into and theflow of effluent out of the Fischer-Tropsch reactor of the otherreaction trains in the plant is continued. The method comprises: flowingthe reactant feed stream to the plurality of reaction trains; dividingthe reactant feed stream into a reactant substream for each reactiontrain; and flowing each reactant substream through a reaction train toconvert the reactants in the reactant substream to a Fischer-Tropschproduct. In an embodiment, the temperature of the Fischer-Tropschproduct flowing out of one reaction train (as measured by the reactorprocess outlet or as measured by the reactor heat exchange fluid outlettemperature) of at least one reactor trains is within about 20° C. orabout 10° C., or about 5° C., or about 2° C., or about 1° C., of thetemperature of the product flowing out of another reaction train withinthe plant.

The Fischer-Tropsch product may be further processed to form alubricating base oil or diesel fuel. For example, the product made inthe microchannel reactor 110 may be hydrocracked and then subjected todistillation and/or catalytic isomerization to provide a lubricatingbase oil, diesel fuel, aviation fuel, and the like. The Fischer-Tropschproduct may be hydroisomerized using the process disclosed in U.S. Pat.No. 6,103,099 or 6,180,575; hydrocracked and hydroisomerized using theprocess disclosed in U.S. Pat. No. 4,943,672 or 6,096,940; dewaxed usingthe process disclosed in U.S. Pat. No. 5,882,505; or hydroisomerized anddewaxed using the process disclosed in U.S. Pat. Nos. 6,013,171,6,080,301 or 6,165,949. These patents are incorporated herein byreference for their disclosures of processes for treatingFischer-Tropsch synthesized hydrocarbons and the resulting products madefrom such processes.

The hydrocracking reaction may be conducted in a hydrocrackingmicrochannel reactor and may involve a reaction between hydrogen and theFischer-Tropsch product flowing from the microchannel reactor 200, orone or more hydrocarbons separated from the Fischer-Tropsch product(e.g., one or more liquid or wax Fischer-Tropsch hydrocarbons). TheFischer-Tropsch product may comprise one or more long chainhydrocarbons. The Fischer-Tropsch product may comprise a Fischer-Tropschwax. In the hydrocracking process, a desired diesel fraction, forexample, may be increased by cracking a C₂₃₊ fraction to mid rangecarbon numbers of C₁₂ to C₂₂. A wax fraction produced from theFischer-Tropsch microchannel reactor 200 may be fed to the hydrocrackingmicrochannel reactor with excess hydrogen for a triple phase reaction.Under reaction conditions at elevated temperatures and pressures, afraction of the liquid feed may convert to a gas phase, while theremaining liquid fraction may flow along the catalyst. In conventionalhydrocracking systems, a liquid stream forms. The use of a microchannelreactor for the hydrocracking reaction enables unique advantages on anumber of fronts. These may include kinetics, pressure drop, heattransfer, and mass transfer.

In an embodiment, the hydrogen, which is in the form of a gas, flows inthe microchannel reactor, and the Fischer-Tropsch product, which is inthe form of a liquid, flows into the gas in the microchannel reactor toform a hydrocracking reactant mixture. This flow pattern facilitatesshearing of the interface between the reactants.

The operating temperature within the hydrocracking microchannel reactormay be in the range of about 200° C. to about 490° C., or about 250° C.to about 450° C., and the temperature of the Fischer-Tropsch synthesisproduct entering the hydrocracking microchannel reactor may be at arelatively low temperature in the range of about 150° C. to about 300°C., or about 150° C. to about 250° C. This relatively low temperature isuseful in preventing coking or clogging in the hydrocarbon microchannelreactor.

The Fischer-Tropsch hydrocarbon products that may be hydrocracked in thehydrocracking microchannel reactor may comprise any hydrocarbon that maybe hydrocracked. These may include hydrocarbons that contain one or moreC—C bonds capable of being broken in a hydrocracking process. TheFischer-Tropsch product that may be hydrocracked may comprise aFischer-Tropsch wax. The hydrocarbons that may be hydrocracked mayinclude saturated aliphatic compounds (e.g., alkanes), unsaturatedaliphatic compounds (e.g., alkenes, alkynes), hydrocarbyl (e.g., alkyl)substituted aromatic compounds, hydrocarbylene (e.g., alkylene)substituted aromatic compounds, and the like.

The feed composition for the hydrocracking microchannel reactor mayinclude one or more diluent materials. Examples of such diluents mayinclude non-reactive hydrocarbon diluents, and the like. The diluentconcentration may be in the range from zero to about 99% by weight basedon the weight of the Fischer-Tropsch product, or from zero to about 75%by weight, or from zero to about 50% by weight. The diluents may be usedto reduce the viscosity of viscous liquid reactants. The viscosity ofthe feed composition in the hydrocracking microchannel reactor may be inthe range from about 0.001 to about 1 centipoise, or from about 0.01 toabout 1 centipoise, or from about 0.1 to about 1 centipoise.

The ratio of hydrogen to Fischer-Tropsch product in the feed compositionentering the hydrocracking microchannel reactor may be in the range fromabout 10 to about 2000 standard cubic centimeters (sccm) of hydrogen percubic centimeter (ccm) of Fischer-Tropsch product, or from about 100 toabout 1800 sccm/ccm, or from about 350 to about 1200 sccm/ccm. Thehydrogen feed may further comprise water, methane, carbon dioxide,carbon monoxide and/or nitrogen.

The H₂ in the hydrogen feed may be derived from another process such asa steam reforming process (product stream with H₂/CO mole ratio of about3), a partial oxidation process (product stream with H₂/CO mole rationof about 2), an autothermal reforming process (product stream with H₂/COmole ratio of about 2.5), a CO₂ reforming process (product stream withH₂/CO mole ratio of about 1), a coal gassification process (productstream with H₂/CO mole ratio of about 1), and combinations thereof. Witheach of these feed streams the H₂ may be separated from the remainingingredients using conventional techniques such as membranes oradsorption.

The hydrocracked Fischer-Tropsch product may comprise a middledistillate fraction boiling in the range of about 260-700° F. (127-371°C.). The term “middle distillate” is intended to include the diesel, jetfuel and kerosene boiling range fractions. The terms “kerosene” and “jetfuel” boiling range are intended to refer to a temperature range of260-550° F. (127-288° C.) and “diesel” boiling range is intended torefer to hydrocarbon boiling points between about 260 to about 700° F.(127-371° C.). The hydrocracked Fischer-Tropsch product may comprise agasoline or naphtha fraction. These may be considered to be the C₅ to400° F. (204° C.) endpoint fractions.

Examples 1-5

A series of synthesis gas interruption tests reported below as Examples1-5 are conducted using a Fischer-Tropsch microchannel reactor. Theresults are reported in Tables 1-5 and 7, and in the attached FIGS.15-20 .

The Fischer-Tropsch microchannel reactor has a process microchannel witha height of 1 mm, a width of 0.6 cm, and a length of 63.5 cm. A 1.9 cmlong SiC bed is positioned in the process microchannel upstream of thecatalyst. The catalyst is a supported cobalt catalyst in the form of afixed bed of particulate solids. The catalyst bed has a length 61.6 cm.Two heat exchange or coolant channels of dimensions 0.2 cm×1.27 cm runparallel to the process channel along its entire length, one coolantchannel on either side of the process microchannel. A hot oil(Marlotherm SH) flow is maintained in both the coolant channels,co-current to the direction of flow of synthesis gas in the processmicrochannel, using a Julabo pump at a minimum flow rate of 8 liters perminute (LPM). The reactor temperature is measured by a set of OmegaK-type thermocouples inserted in thermowells in the metal webs betweenthe process microchannel and coolant channels.

The packed apparent bed density (PABD) of the SiC and the catalyst ismeasured ex-situ by measuring the mass of SiC and catalyst filled in aClass-A 10 ml “to contain” graduated cylinder and densified using aQuantachrome Autotap set to 1500 taps. After installing a retentionassembly at the reactor outlet, the catalyst is first loaded to a bedlength of 61.6 cm followed by SiC to top off the process microchannel.The bed is densified. The PABD of the catalyst and SiC in the processmicrochannel is within about ±5% of the ex-situ measured PABD. This isfollowed by installation of a catalyst retention assembly at the reactorinlet.

The reactor is installed in a test stand and connections are providedfor feeding synthesis gas into the process microchannel and flowingeffluent out of the process microchannel. Similarly, connections areprovided for feeding hot oil into and out of the heat exchange channels.The catalyst is activated using hydrogen with a gas hourly spacevelocity (GHSV) of 7000 hr⁻¹ based on the loaded volume of catalyst inthe reactor. The temperature is ramped from ambient to an activationtemperature of 350° C. After the completion of an activation hold, thereactor is cooled to a temperature of 150-170° C. and a synthesis gasfeed is introduced to the reactor. The reactor pressure is adjusted tothe target value and the temperature is ramped up to achieve the desiredCO conversion. The reactor performance (CO conversion, CH₄ selectivity,etc) is monitored by measuring the tail gas flow and effluentcomposition from the reactor outlet.

The bottling procedure may be performed using the apparatus illustratedin FIG. 21 . Referring to FIG. 21 , before initiating the bottlingsequence, the pressure related (PT) interlocks are disabled and thepressure control valve (PCV) is set in manual mode with output equal tothe current operating value in automatic mode. The ball valves (BV) atthe reactor outlet and reactor inlet are closed (in that sequence) toisolate the synthesis gas within the reactor. The three-way BV after thePCV is also closed to prevent depressurization of the product tanks.Finally, all gas flows to the reactor are turned off by closing the BVsafter the mass flow controllers (MFCs) and then adjusting the MFC outputto fully closed.

If depressurization is desired, the PCV output is set to the same valueas it was prior to bottling the reactor. Then, the BV at the inlet ofreactor is opened, the BV at the outlet of the reactor is opened, andthe three-way BV after the PCV is opened (in that sequence). The PCV isset to ramp to desired rate to depressurize the reactor and system tothe target pressure.

If purge is required, the mass flow controller (MFC) BV for the desiredpurge gas (e.g., hydrogen, natural gas, nitrogen) is opened and flowsturned ON at the target flow rate and the gas is allowed to flow for thedesired time. After the completion of the purge, the reactor is againbottled as described above.

If holding at temperature, no changes in reactor temperature setting aremade. If cooling down, the reactor temperature is changed at the desiredrate to the target value by setting a ramp rate for the average reactortemperature control.

The restart of operations from a bottled scenario can be achieved in thefollowing manner: To unbottle the reactor, the inlet BV to the reactoris opened followed by the outlet BV to the reactor and the three-way BVafter the PCV. Synthesis gas flow is restarted by setting the desiredN₂, H₂, and CO flows, resulting in establishment of full flows within afew seconds. The reactor is then pressurized to the desired inletpressure using the PCV. If needed, the reactor is heated to the targettemperature by setting a ramp rate for the average reactor temperaturecontrol.

The reaction rate model for Fischer Tropsch synthesis on cobalt catalyst(Yates and Satterfield, 1991, referred to above) may be given by

${- R_{CO}} = \frac{a\mu_{CO}\mu_{H_{2}}}{\left( {1 + {bP}_{CO}} \right)^{2}}$where a and b are temperature-dependent constants; “a” representing akinetic parameter and “b” an adsorption coefficient as shown below:

Reactor Temperature (° C.)$a\left( \frac{mmol}{\min g_{cat}{MPa}^{2}} \right)$$b\left( \frac{1}{MPa} \right)$ 240 75.76 11.61 220 53.11 22.26Assuming: a=Ae^(−E) ^(a) ^(/RT) and b=Be^(−H/RT)The kinetic rate equation parameters are estimated as,

${a = {0.082505e^{{{- 37367}/8.314}T}{in}{\frac{\frac{mols}{g}}{g_{cat}}/{atm}^{2}}{and}}}{b = {1.259 \times 10^{- 7}e^{{69475/8.314}T}{in}{1/{{atm}.}}}}$

Example 1

Example 1 is provided for purposes of comparison. The test reported inTable 1 and FIG. 15 , is conducted using a nitrogen purge followed by afast reactor cool down to 170° C. Prior to the start of the stoppage,the microchannel reactor is operating under Fischer-Tropsch conditionswith a synthesis gas feed having a H₂:CO molar ratio of 1.77, 32% byvolume inerts with an inlet pressure of 350 psig (2413 kilopascals), anda contact time of 290 milliseconds. The average reactor operatingtemperature is 215° C. To perform synthesis stoppage, the synthesis gasfeed is stopped and a nitrogen purge is introduced with a flowequivalent to a 1.6 second contact time for 3 minutes. The microchannelreactor is bottled and cooled to a temperature of 170° C. at a rate of30° C. per hour. The temperature is held at 170° C. for 20 hours. Thefull flow of synthesis gas (290 milliseconds contact time) into themicrochannel and the flow of effluent out of the reactor is thenstarted. Once pre-synthesis stop conditions with feed gas at a H₂:COmolar ratio of 1.77 with 32% by volume inerts, inlet pressure of 350psig (2413 kilopascals), and a contact time of 290 milliseconds areattained, the microchannel reactor is heated to the target(pre-stoppage) temperature at a rate of 15° C. per hour. This Examplemay be considered as representative of the prior art (with the exceptionof the temperature ramp rates used for reactor heat-up), and is providedfor purposes of comparison. The results for this Example show a COconversion loss of about 1.5% which would be equivalent to an activitytemperature delta of 1.3° C. for the process. As calculated in Table 7,the relative activity ratio for this Example is 0.979.

TABLE 1 24 his avg before interrupt 12-24 his avg after restart COconversion 71.7% 70.2% CH₄ selectivity  7.4%  7.1% CO₂ selectivity  0.5% 0.5% C₂ selectivity  0.8%  1.0% C₃ selectivity  2.3%  2.0% C₄selectivity  2.8%  3.0% C₅+ selectivity 86.2% 85.9%

Example 2

The test reported in Table 2 and FIG. 16 , is conducted with themicrochannel reactor bottled and simultaneously cooled down to 185° C.Prior to performing the process stoppage, the microchannel reactor isoperating under Fischer-Tropsch conditions with a synthesis gas feedhaving H₂:CO molar ratio of 1.79, 28% by volume inerts with an inletpressure of 350 psig (2413 kilopascals), and a contact time of 310milliseconds. The average reactor operating temperature is 215° C. Toperform synthesis stoppage, the synthesis gas feed is stopped and theflow of effluent out of the reactor is stopped and the reactor is cooledto 185° C. at a rate of 30° C. per hour, and then held at 185° C. for 1hour. The flow of synthesis gas into the microchannel reactor and theflow of effluent out of the reactor is then started. The pre-synthesisstop conditions with feed gas at a H₂:CO molar ratio of 1.79 with 28% byvolume inerts, inlet pressure of 350 psig (2413 kilopascals), and acontact time of 310 milliseconds are re-established within about 10-30seconds. The microchannel reactor is heated to the target (pre-stoppage)temperature at a rate of 15° C. per hour. These tests show a similarlevel of CO conversion loss as the nitrogen purge described incomparative Example 1 which would be an activity temperature delta ofabout 1.3° C. As calculated in Table 7, the relative activity ratio forthis Example is also 0.979. This procedure has the effect of eliminatingthe purge requirement and enables restoring the pre-stoppage conditionsfaster. This improves the availability of the reactor for synthesisoperation relative to that in Example 1.

TABLE 2 24 his avg before interrupt 12-24 his avg after restart COconversion 72.6% 71.1  CH₄ selectivity  6.6%   6.9% CO₂ selectivity 0.4%   0.4% C₂ selectivity  0.8%   1.0% C₃ selectivity  2.3%   2.7% C₄selectivity  2.7%   3.0% C₅+ selectivity 87.1% 86.1%

Example 3

The test reported in Table 3 and FIG. 17 , is conducted with themicrochannel reactor bottled up at the reaction temperature of 215° C.Prior to performing the process stoppage, the microchannel reactor isoperating under Fischer-Tropsch conditions with a synthesis gas feedhaving a H₂:CO molar ratio of 1.79, 28% by volume inerts with an inletpressure of 350 psig (2413 kilopascals), and a contact time of 310milliseconds. The average reactor operating temperature is 215° C. Tosimulate synthesis stoppage, the synthesis gas feed is stopped and theflow of effluent out of the reactor is stopped. The reactor temperatureis maintained within ±1° C. of 215° C. during the bottling. At the endof a bottled period of 1 hour, the process is restarted with synthesisgas flowing into the reactor and effluent flowing out. The pre-synthesisstop conditions with feed gas at a H₂:CO molar ratio of 1.79 with 28% byvolume inerts, inlet pressure of 350 psig (2413 kilopascals), and acontact time of 310 milliseconds are re-established within 10-30seconds. There is a CO conversion loss of about 0.6% which represents a50% improvement over the results shown in comparative Example 1. Thiscorresponds to an activity temperature delta of 0.5° C. As calculated inTable 7, the relative activity ratio for this Example is 0.992. Withthis procedure, not only is the loss of catalyst activity significantlyreduced, but the process also allows the reactor to be restarted byopening the reactor inlet and outlet to start flowing synthesis gas intothe reactor and effluent out of the reactor in a matter of minutes, asthe reactor is already at the target operating conditions. Given theconcerns of carbon deposition (one of the primary causes of catalystdeactivation) on the catalyst under low hydrogen partial pressures (asH₂ will react to extinction), the decrease in loss of catalyst activitycompared to the prior art shown in Example 1 is an unexpected andbeneficial result.

TABLE 3 24 his avg before interrupt 12-24 his avg after restart COconversion 70.8% 70.2% CH₄ selectivity  6.8%  6.8% CO₂ selectivity  0.4% 0.4% C₂ selectivity  0.9%  1.0% C₃ selectivity  2.6%  2.8% C₄selectivity  3.0%  3.1% C₅+ selectivity 86.3% 85.9%

Example 4

The test reported in Table 4 and FIG. 18 , is conducted with themicrochannel reactor bottled, followed by a low pressure nitrogen purgeand slow cool down to 170° C. The process stoppage in this case wouldresemble a loss of coolant scenario where the initial response would bethat of the reactor is bottled to be later followed by a nitrogen purgeand slow cooling representative of the drop in reactor temperature withheat losses. Prior to performing the process stoppage, the microchannelreactor is operating under Fischer-Tropsch conditions with a synthesisgas feed having a H₂:CO molar ratio of 1.79, 28% by volume inerts withan inlet pressure of 350 psig (2413 kilopascals), and a contact time of310 milliseconds. The average reactor operating temperature is 215° C.To simulate synthesis stoppage, the synthesis gas feed is stopped andthe reactor is bottled for a period of 1 hour. At the end of 1 hour ofbottled time, the pressure within the reactor is reduced from 350 psig(2413 kilopascals) to 35 psig (241.3 kilopascals) at a rate of 500 psi(3447 kilopascals) per hour. The microchannel reactor is then purgedwith N₂ at 125.8 standard cubic centimeters per minute (sccm) for 3minutes. The microchannel reactor is then bottled for 24 hours under N₂with a slow cooldown to 170° C. at a rate of 2° C. per hour. After 24hours, the flow of synthesis gas into the microchannel reactor and theflow of effluent out of the microchannel reactor is then restarted. Thepre-synthesis stop conditions with feed gas at a H₂:CO molar ratio of1.79 with 28% by volume inerts and a contact time of 310 millisecondsare re-established within 10-30 seconds. The microchannel reactor isthen pressurized to 350 psig (2413 kilopascals) and subsequently heatedto the target (pre-stoppage) operating temperature at a rate of 15° C.per hour. The CO conversion loss after restart is 1.3%, which would beequivalent to an increase in the reactor temperature of 1.1° C., thatis, an activity temperature delta of about 1° C. As calculated in Table7, the relative activity ratio for this Example is 0.981.

TABLE 4 24 his avg before interrupt 12-24 his avg after restart COconversion 69.1% 67.8% CH₄ selectivity  6.8%  6.7% CO₂ selectivity  0.4% 0.3% C₂ selectivity  0.9%  1.0% C₃ selectivity  2.5%  2.7% C₄selectivity  2.9%  3.0% C₅+ selectivity 86.5% 86.2%

Example 5

The test reported in Table 5 and FIG. 19 is conducted with themicrochannel reactor bottled for 1 hour and then slowly cooled down to175° C. Prior to performing the process stoppage, the microchannelreactor is operating under Fischer-Tropsch conditions with a syngas feedhaving a H₂:CO molar ratio of 1.79, 28% by volume inerts with an inletpressure of 350 psig (2413 kilopascals), and a contact time of 310milliseconds. The average reactor operating temperature is 223° C. Tosimulate synthesis stoppage, the synthesis gas feed is stopped and thereactor is bottled for a period of 1 hour. At the end of 1 hour ofbottling time, the microchannel reactor is cooled down to 175° C. at arate of 2° C. per hour. After 24 hours, the flow of synthesis gas intothe reactor and the flow of effluent out of the reactor are started. Thepre-synthesis stop conditions with feed gas at a H₂:CO molar ratio of1.79 with 28% by volume inerts, inlet pressure of 350 psig (2413kilopascals), and a contact time of 310 milliseconds are re-establishedwithin 10-30 seconds. The microchannel reactor is heated to thepre-stoppage reaction temperature at a rate of 15° C. per hour. This isa modification of the test shown in Example 4 to eliminate the use of aninert N₂ purge to achieve a similar loss in CO conversion by merebottling of the reactor under operating synthesis gas. The bottledreactor under synthesis gas is shown to have a CO conversion loss ofabout 2.1%, which would be equivalent to an activity temperature deltaof 1.7° C. As calculated in Table 7, the relative activity ratio forthis Example is 0.971.

TABLE 5 24 his avg before interrupt 12-24 his avg after restart COconversion 71.8% 69.7% CH₄ selectivity  8.6%  9.0% CO₂ selectivity  0.7% 0.7% C₂ selectivity  1.3%  1.4% C₃ selectivity  3.1%  3.5% C₄selectivity  3.6%  4.0% C₅+ selectivity 82.6% 81.4%

Example 6

A series of synthesis gas interruption tests are conducted using amulti-channel Fischer-Tropsch microchannel reactor. The results arereported in Table 6 and FIG. 20 . The microchannel reactor has 850process microchannels (arranged in two layers) with dimensions similarto the single channel described above for Examples 1-5. A supportedcobalt catalyst in the form of a fixed bed of particulate solids isloaded in the process microchannels with a bed length of 40 cm. Each ofthe process layers has an adjacent coolant layer. Each coolant layer has128 coolant microchannels. Partial boiling heat transfer is used toremove the reaction heat by flowing water in the coolant circuits.Headers and footers are attached for the coolant and processmicrochannels to have external connections to macroscale (i.e., larger),piping.

The Fischer-Tropsch synthesis process conditions in the reactor areinterrupted, and then a hydrogen treatment is performed at a temperatureof 350-375° C. Following this treatment the reactor is cooled to atemperature of 70° C. and 100% air (21% O₂) is introduced into thereactor. The exotherm is controlled by adjusting one or more of the airflow, coolant flow, and coolant temperature. It is preferred to maintainthe exotherm during the air treatment to less than a 15° C. temperaturerise, or less than 10° C., or less than 5° C. The catalyst is oxidizedduring the air treatment. Cobalt oxidation with air is a highlyexothermic reaction. Despite this only a minor temperature increase ofabout 7° C. is observed in this example in the coolant outlettemperature due to the excellent heat transfer characteristics of themicrochannel reactor. This is well within a level that can be toleratedby the reactor coolant system without a negative impact on the catalystactivity. This result enables the use of low pressure air for performingcatalyst oxidation using simplified equipment provided that the flow ofair is controlled in order to avoid excessive catalyst temperatures. Italso eliminates the need for blending of the air with a N₂ stream todilute the oxygen concentration, which in turn significantly reduces theneed for nitrogen availability at the plant site.

Following a final reduction treatment in a flowing hydrogen environmentat a temperature of about 350° C., the catalyst is brought to the targetoperating conditions in 18 hours. Because of the oxidative regeneration,the catalyst activity is restored to a level higher than that before theinterruption of the Fischer-Tropsch synthesis process. The H₂:CO molarratio is 1.83, with 41% by volume inerts. The inlet pressure is 350 psig(2413 kilopascals). The contact time is 355 to 360 milliseconds. Usingthis regenerative procedure, the activity temperature delta is reducedto −16° C. In the Table 6 below, all percentages are by volume. Ascalculated in Table 7, the relative activity ratio for this Example is4.036.

TABLE 6 24 his avg before interrupt 12-24 his avg after restart Avg.reactor 218° C. 202° C. temperature CO conversion 69.3% 70.9% CH₄selectivity 10.4%  9.2% C₂ selectivity  1.6%  1.0% C₃ selectivity  3.3% 2.3% C₄ selectivity  3.2%  2.7% C₅+ selectivity 80.1% 84.4%

Table 7 below summarizes the calculation of relative activity ratio forthe Examples 1-6 using the process performance data and the intrinsickinetic model from Yates & Satterfield as described above.

TABLE 7 Inlet Pressure Partial pressure Partial pressure Contact CO flowto reactor Example (psig) H2:CO inerts H2 (atm) CO (atm) time (ms) GHSV(1/h) (v/vcat/h) 1 350 1.77 32% 10.78 6.09 290 12414 3047 2 350 1.79 28%11.46 6.40 310 11613 2997 3 350 1.79 28% 11.46 6.40 310 11613 2997 4 3501.79 28% 11.46 6.40 310 11613 2997 5 350 1.79 28% 11.46 6.40 310 116132997 6 350 1.83 41%  9.47 5.17 358 10056 2096 before after activityactivity relative activity (CO (CO ratio CO reacted YS rate reacted/ COreacted YS rate reacted/ (activity-after/ Example CO conv ART (C)v/vcat/h v/vcat/h YS rate) CO conv ART (C) v/vcat/h v/vcat/h YS rate)activity-before) 1 71.7% 215.0 2185.0 146.5 14.9 70.2% 215.0 2139.3146.5 14.6 0.979 2 72.6% 215.0 2175.7 149.0 14.6 71.1% 215.0 2130.8149.0 14.3 0.979 3 70.8% 215.0 2121.8 149.0 14.2 70.2% 215.0 2103.8149.0 14.1 0.992 4 69.1% 215.0 2070.8 149.0 13.9 67.8% 215.0 2031.9149.0 13.6 0.981 5 71.8% 223.0 2151.8 287.9  7.5 69.7% 223.0 2088.8287.9  7.3 0.971 6 69.3% 218.0 1452.8 190.2  7.6 70.9% 202.0 1486.4 48.2 30.8 4.036

While the invention has been explained in relation to variousembodiments, it is to be understood that various modifications thereofwill become apparent to those skilled in the art upon reading thespecification. Therefore, it is to be understood that the inventiondisclosed herein includes any such modifications that may fall withinthe scope of the appended claims.

The invention claimed is:
 1. A method for restarting a synthesis gasconversion process, wherein the synthesis gas conversion processcomprises flowing synthesis gas into a reactor in contact with asynthesis gas conversion catalyst at a desired reaction temperature andpressure to produce a synthesis gas conversion product and flowingeffluent comprising the synthesis gas conversion product out of thereactor, the method comprising: (A) stopping the flow of synthesis gasinto the reactor; (B) flowing natural gas into the reactor to purge thereactor; and (C) restarting the flow of synthesis gas into the reactor;wherein the reactor comprises a microchannel reactor.
 2. The method ofclaim 1 wherein the synthesis gas comprises CO and prior to stopping theflow of synthesis gas into the reactor the conversion of CO is at afirst conversion value, and after restarting the flow of synthesis gasinto the reactor the conversion of CO at the first conversion value isachieved within a time period of up to about 3 hours.
 3. The method ofclaim 1 wherein the reactor comprises a fixed bed reactor, a fluidizedbed reactor or a slurry phase reactor.
 4. The method of claim 1 whereinthe synthesis gas conversion process comprises a process for convertingsynthesis gas to methane.
 5. The method of claim 1 wherein the synthesisgas conversion process comprises a process for converting synthesis gasto methanol or dimethyl ether.
 6. The method of claim 1 wherein thesynthesis gas conversion process is a Fischer-Tropsch process.
 7. Themethod of claim 6 wherein the synthesis gas catalyst is aFischer-Tropsch catalyst and the synthesis gas conversion product is aFischer-Tropsch product.
 8. The method of claim 6 wherein a tail gas isproduced in the reactor and at least part of the tail gas is combinedwith the synthesis gas to form a reactant mixture, the volumetric ratioof the synthesis gas to tail gas being in the range from about 1:1 toabout 10:1.
 9. The method of claim 6 wherein the synthesis gas comprisesH₂ and CO, the mole ratio of H₂ to CO being in the range from about1.4:1 to about 2.1:1.
 10. The method of claim 6 wherein the synthesisgas comprises fresh synthesis gas, the fresh synthesis gas comprisingCO, the conversion of CO from the fresh synthesis gas being at leastabout 70%.
 11. The method of claim 6 wherein the selectivity to methanein the Fischer-Tropsch product is in the range from about 0.01 to 10%.12. The method of claim 7 wherein the Fischer-Tropsch catalyst comprisescobalt and a support.
 13. The method of claim 12 wherein theFischer-Tropsch catalyst has a cobalt loading in the range from about 10to about 60% by weight.
 14. The method of claim 12 wherein the supportcomprises a refractory metal oxide, carbide, carbon, nitride, or amixture of two or more thereof.
 15. The method of claim 12 wherein thesupport comprises alumina, zirconia, silica, titania, or a mixture oftwo or more thereof.
 16. The method of claim 12 wherein the surface ofthe support is modified by being treated with titania, zirconia,magnesia, chromia, alumina, or a mixture of two or more thereof.
 17. Themethod of claim 12 wherein the support comprises silica and the surfaceof the support is modified by being treated with titania.
 18. The methodof claim 1 wherein the microchannel reactor comprises one or moremicrochannel reactor cores, each microchannel reactor core comprisingone or more layers of process microchannels and one or more layers ofheat exchange channels.
 19. The method of claim 18 wherein themicrochannel reactor comprises a plurality of the layers of processmicrochannels and a plurality of the layers of heat exchange channels, acatalyst being in the process microchannels, each layer of heat exchangechannels being in thermal contact with at least one layer of processmicrochannels, a header for flowing reactants into the processmicrochannels, a footer for receiving product flowing out of the processmicrochannels, a header for flowing a heat exchange fluid into the heatexchange channels, and a footer for receiving heat exchange fluidflowing out of the heat exchange channels.
 20. The method of claim 18wherein a plurality of the microchannel reactor cores are positioned ina containment vessel, each microchannel reactor core comprising aplurality of layers of process microchannels and a plurality of layersof heat exchange channels, a catalyst being in the processmicrochannels, each layer of heat exchange channel being in thermalcontact with at least one layer of process microchannels, thecontainment vessel being equipped with a manifold for flowing thereactants to the microchannel reactor cores, a manifold for flowingproduct from the microchannel reactor cores, a manifold for flowing aheat exchange fluid to the microchannel reactor cores, and a manifoldfor flowing heat exchange fluid from the microchannel reactor cores. 21.The method of claim 20 wherein the containment vessel contains from 1 toabout 12 microchannel reactor cores.
 22. The method of claim 18 whereineach layer of process microchannels has a height of up to about 10 mm.23. The method of claim 22 wherein each layer of process microchannelshas a length of up to about 5 meters and a width of up to about 5meters.
 24. The method of claim 18 wherein each microchannel reactorcore is made of a material comprising: aluminum; titanium; nickel;copper; an alloy of any of the foregoing metals; steel; stainless steel;monel; inconel; brass; quartz; silicon; or a combination of two or morethereof.
 25. The method of claim 18 wherein surface features are presentin the process microchannels.
 26. The method of claim 18 wherein theheat exchange channels are microchannels.
 27. The method of claim 1wherein the catalyst is in the form of a fixed bed of particulatesolids.
 28. The method of claim 1 wherein the catalyst is aFischer-Tropsch catalyst is in the form of a fixed bed of particulatesolids, the particulate solids having a median particle diameter in therange from about 1 to about 1000 microns.
 29. The method of claim 1wherein the catalyst is positioned in a process microchannel and iscoated on interior walls of the process microchannels or grown oninterior walls of the process microchannel.
 30. The method of claim 1wherein the catalyst is supported on a support having a flow-byconfiguration, a flow-through configuration, or a serpentineconfiguration.
 31. The method of claim 1 wherein the catalyst issupported on a support having the configuration of a foam, felt, wad,fin, or a combination of two or more thereof.
 32. The method of claim 1wherein the catalyst is supported on a support in the form of a finassembly.
 33. The method of claim 1 wherein the catalyst is supported ona corrugated insert.
 34. The method of claim 1 wherein the reactiontemperature is in the range from about 150 to about 300° C.
 35. Themethod of claim 1 wherein the contact time of reactants with thecatalyst is up to about 3600 milliseconds.
 36. The method of claim 1wherein the product is a Fischer-Tropsch product which compriseshydrocarbons boiling at a temperature of at least about 30° C. atatmospheric pressure.
 37. The method of claim 1 wherein the product is aFischer-Tropsch product which comprises hydrocarbons boiling above atemperature of about 175° C. at atmospheric pressure.
 38. The method ofclaim 1 wherein the product is a Fischer-Tropsch product which comprisesparaffins and/or olefins of 2 to about 200 carbon atoms.
 39. The methodof claim 1 wherein the product is a Fischer-Tropsch product which isfurther processed using separation, fractionation, hydrocracking,hydroisomerizing, dewaxing, or a combination of two or more thereof. 40.The method of claim 1 wherein the product is a Fischer-Tropsch productwhich is further processed to form an oil of lubricating viscosity or amiddle distillate fuel.
 41. The method of claim 18 wherein the layer ofprocess microchannels has fluid flowing in it in one direction, and thelayer of heat exchange channels has fluid flow in a direction that isco-current or counter-current to the flow of fluid in the layer ofprocess microchannels.
 42. The method of claim 18 wherein the layer ofprocess microchannels has fluid flowing in it in one direction, and thelayer of heat exchange channels has fluid flowing in it in a directionthat is cross-current to the flow of fluid in the process microchannels.43. The method of claim 1 wherein the reactor comprises a processmicrochannel and after restarting the flow of synthesis gas into thereactor, the synthesis gas flows in the process microchannel at asuperficial velocity of at least about 0.01 m/s.
 44. The method of claim1 wherein the reactor comprises a process microchannel and afterrestarting the flow of synthesis gas into the reactor, the synthesis gasflows in the process microchannel at a space velocity of at least about1000 hr⁻¹.
 45. The method of claim 1 wherein the microchannel reactor isa microchannel reactor with layers of process microchannels formed bypositioning a waveform between planar sheets.
 46. The method of claim 1wherein the microchannel reactor comprises a plurality of plates in astack defining a plurality of process layers and a plurality of heatexchange layers, each plate having a peripheral edge, the peripheraledge of each plate being welded to the peripheral edge of the nextadjacent plate to provide a perimeter seal for the stack.
 47. The methodof claim 1 wherein the synthesis gas comprises CO and H₂ and thedeactivation rate of the catalyst is less than a loss of about 0.2% COconversion per day.
 48. The method of claim 1 wherein the reactiontemperature in the reactor is at a first reaction temperature and theconversion of CO is at a first conversion value prior to stopping theflow of synthesis gas into the reactor, and after restarting the flow ofsynthesis gas into the reactor the conversion of CO in the reactor is atthe first conversion value and the temperature in the reactor is at asecond reaction temperature, the second reaction temperature being lessthan about 5° C. higher than the first reaction temperature.
 49. Themethod of claim 1 wherein the natural gas flowing into the reactorduring step (B) is a desulfurized natural gas.
 50. The method of claim 1wherein prior to the step of restarting the flow of synthesis gas intothe reactor, the temperature within the reactor is below the desiredreaction temperature, the temperature of the reactor being increased tothe desired reaction temperature at a rate of up to about 5° C. perhour.
 51. The method of claim 1 wherein prior to stopping the flow ofsynthesis gas into the reactor, the temperature in the reactor is at adesired operating temperature, and during the period of time betweenstopping the flow of synthesis gas into the reactor and restarting theflow of synthesis gas into the reactor the temperature in the reactor iswithin about 20° C. of the desired operating temperature.
 52. The methodof claim 1 wherein the activity temperature delta for the process afterrestarting the flow of synthesis gas into the reactor is up to about 5°C.
 53. The method of claim 1 wherein the relative activity ratio for theprocess is at least about 0.85.
 54. The method of claim 1 wherein thereactor includes a heat transfer surface, and the ratio of the surfacearea of the heat transfer surface to the volume of catalyst in thereactor is at least about 300 m² of heat transfer surface area per m³ ofcatalyst.
 55. The method of claim 1 wherein the temperature in thereactor is at a pre-stoppage temperature prior to stopping the flow ofsynthesis gas into the reactor, and the temperature of the reactorduring the step of restarting the flow of synthesis gas into the reactoris at the pre-stoppage temperature or below the pre-stoppage temperatureby up to about 5° C.